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Dissertation

Biomass upgrading technologiesfor carbon-neutral and carbon-negative

electricity generationTechno-economic analysis of hydrothermal carbonizationand comparison with wood pelletizing, torrefaction and

anaerobic digestion

Dipl.-Ing. Berit Erlach

Berlin 2014

Technische Universität BerlinInstitute for Energy Engineering

Biomass upgrading technologies for carbon-neutral andcarbon-negative electricity generation

Techno-economic analysis of hydrothermal carbonization andcomparison with wood pelletizing, torrefaction and anaerobic digestion

vorgelegt vonDiplom-Ingenieurin

Berit Erlachgeboren in Uppsala (Schweden)

von der Fakultät III – Prozesswissenschaftender Technischen Universität Berlin

zur Erlangung des akademischen Grades

Doktor der Ingenieurwissenschaften– Dr.-Ing. –

genehmigte Dissertation

Promotionsausschuss:

Vorsitzender: Prof. Dr.-Ing. Felix Ziegler

Gutacher: Prof. Dr.-Ing. Georgios Tsatsaronis

Gutachter: : Prof. em. Mats Westermark

Tag der wissenschaftlichen Aussprache: 04. Juli 2014

Berlin 2014D 83

Danksagung

Die vorliegende Arbeit entstand während meiner Tätigkeit als wissenschaftliche Mitarbei-terin am Fachgebiet Energietechnik und Umweltschutz des Instituts für Energietechnikan der Technischen Universität Berlin. Von 2009–2012 bestand eine Förderung durch dasBundesministerium für Bildung und Forschung im Rahmen des Verbundprojektes “Hy-drothermale Karbonisierung von Biomasse — Potential, Entwurf, Pilotanlage”, Förder-kennzeichen 01LS0806B.

Ich bedanke mich bei Herrn Prof. George Tsatsaronis für die Betreuung und Begutachtungder Dissertation und bei Herrn Prof. Mats Westermark für die Übernahme des Korreferats.

Bei Jan Stemann, Axel Funke, Benjamin Wirth, Uwe Kracht und all den anderen Projekt-beteiligten möchte ich mich für die intensive und konstruktive Zusammenarbeit bei derErforschung der HTC bedanken. Den Studierenden, deren Abschlussarbeiten ich betreuthabe, danke ich für viele inspirierende Diskussionen und wertvolle Anregungen für meineArbeit.

Bei den Kollegen am Fachgebiet bedanke ich mich für die freundschaftliche Zusammen-arbeit und die schöne Zeit.

Zu guter Letzt bedanke ich mich bei Robbie Morrison für seine Unterstützung, seinekritischen Fragen und Anregungen, und das mehrfache Korrekturlesen und “brushing upof the english”.

This work is licensed under the Creative Commons Attribution-NonCommercial-ShareAlike4.0 International License. To view a copy of this license, visit

http://creativecommons.org/licenses/by-nc-sa/4.0/.

Zusammenfassung

Biomasse hat sich in den letzten Jahren zu einem international gehandelten Brennstoffentwickelt. Aufbereiteten Biobrennstoffen wie Holzpellets oder torrefiziertem Holz kommtvor diesem Hintergrund eine besondere Bedeutung zu, da sie für Transport und Lagerungbesser geeignet sind als unbehandelte Biomasse. Die hydrothermale Karbonisierung (HTC)ist ein Verfahren, welches die Herstellung eines kohleähnlichen Festbrennstoffs aus Biomas-se ermöglicht. In dieser Arbeit wird die HTC unter technisch-wirtschaftlichen Aspektenuntersucht und mit konkurrierenden Verfahren der Biomasseaufbereitung verglichen.Im ersten Teil der Arbeit werden HTC, Holzpelletierung, Torrefizierung und Biogaser-zeugung anhand von Energieeffizienz, Treibhausgasemissionen und Produktionskosten be-wertet. Der betrachtete Einsatzfall der Biobrennstoffe ist dabei der Ersatz von fossilenBrennstoffen in Bestandskraftwerken. Im zweiten Teil wird die Bedeutung der Aufberei-tungsverfahren für die Biomasseverstromung mit CO2-Abtrennung (BECCS) untersucht.Torrefizierung und HTC zerstören die faserige Struktur der Biomasse und ermöglichen soihren Einsatz in Flugstromvergasern. Torrefizierung oder HTC gefolgt von Flugstromver-gasung wird verglichen mit der Wirbelschichtvergasung von unbehandeltem Holz.Die untersuchten Prozesse werden mit dem Simulationsprogramm Aspen Plus modelliertund anhand von Exergieanalyse thermodynamisch bewertet. Für die HTC werden darüberhinaus mit Hilfe einer exergoükonomischen Analyse Potenziale zur Reduktion der Produk-tionskosten identifiziert. Die Transportkosten für die Biomasse werden mit einem einfachenModell für die Lieferkette in Abhängigkeit von der Anlagenkapazität abgeschätzt.Die Ergebnisse zeigen, dass HTC nur dann mit Holzpelletierung konkurrenzfähig ist, wennbiogene Abfälle als Einsatzstoff verwendet werden. Hierfür müssen ganzjährig große Abfall-mengen, um die 100 kt/a, zur Verfügung stehen. Geeignete Einsatzstoffe sind kommunaleGrünabfälle und leere Fruchstände aus der Palmölproduktion.Die Exergienanalyse zeigt, dass bei allen Verfahren zur Herstellung fester Biobrennstoffedie Trocknung der Rohbiomasse oder des aufbereiteten Brennstoffs für eine großen Teilder Exergievernichtung verantwortlich ist. Bei HTC und Biogaserzeugung treten außer-dem hohe Exergieverluste durch Abwasser bzw. Gärrest auf. Potenzielle Maßnahmen zurEffizienzverbesserung und Kostenreduktion bei der HTC beinhalten eine effiziente Wärme-rückgewinnung, Trocknung in Dampfatmosphäre, und Biogasgewinnung aus dem Abgas.Darüber hinaus könnte die wärmeseitige Integration des HTC-Prozesses in ein Dampf-kraftwerk die Vorwärmung der Biomasse vereinfachen und die Kosten reduzieren.Beim Einsatz in Kraftwerken mit Biomassevergasung und CO2-Abscheidung bietet Flug-stromvergasung von torrefiziertem Holz oder HTC-Kohle einen höheren Wirkungsgradgegenüber der Wirbelschichtvergasung von unbehandeltem Holz. Allerdings kann der Ef-fizienzgewinn im Kraftwerksprozess die Umwandlungsverluste der Biomasseaufbereitungnicht kompensieren. Zudem ist die CO2-Abscheiderate bezogen auf den Kohlenstoff derRohbiomasse bei dem Konversionspfad über Torrefizierung oder HTC mit 66–69% deutlichniedriger als bei der direkten Vergasung der Biomasse mit 82–86%. Ist der CO2-Preis aus-reichend hoch, um CO2-Abscheidung in fossil gefeuerten Kraftwerken rentabel zu machen,so sind auch einige der untersuchten BECCS-Konzepte nahezu konkurrenzfähig.

Abstract

Biomass is increasingly becoming an internationally traded commodity fuel. In this con-text, biomass upgrading technologies such as torrefaction and hydrothermal carbonization(HTC), which increase the energy density and improve the storability and handling of thebiomass, have recently gained attention. This work provides a techno-economic assessmentof commercial-scale HTC plants and their competing technologies.In the first part of this work, energetic efficiency, GHG emissions and costs of HTC arecompared to those of wood pelletizing, torrefaction and anaerobic digestion. Moreover,the substitution of fossil fuels by the aforementioned biofuels in existing power stations isanalyzed. The second part focusses on the potential role of biomass upgrading technologiesfor bioenergy with carbon capture and storage (BECCS). Torrefaction and HTC causethe biomass to lose its fibrous structure, thus facilitating entrained flow gasification. Inorder to investigate the merits of this conversion pathway, torrefaction or HTC followedby entrained flow gasification is compared to the direct fluidized bed gasification of rawwood.The analysis is based on flowsheet simulations created with Aspen Plus. Exergy analysis isemployed to locate thermodynamic losses within the respective processes. Exergoeconomicanalysis is applied to the HTC plant design to reveal potentials for reducing the biocoalproduction costs. A simple model of the entire supply chain is developed in order to assessthe costs and GHG emissions related to biomass and biofuel transport and storage andtheir dependency on the plant capacity.The results indicate that HTC can only be economically competitive with conventionalwood pelletizing if waste biomass is used as a feedstock. Depending on the remunerationfor waste disposal, relatively large processing capacities of up to 100 kt/a of feedstock arerequired year-round to make HTC an economically viable proposition. Potential feedstocksinclude park and gardening waste and empty fruit bunches from palm oil production.Exergy analysis reveals that drying of the feedstock or biofuel is the most significant sourceof exergy destruction in all the analyzed processes generating solid biofuels. HTC and an-aerobic digestion also suffer large exergy losses through their waste streams. Measures toimprove the efficiency and cost of HTC include efficient heat recovery, drying in super-heated steam, and using the waste water to produce biogas. Integration of HTC with arankine-cycle CHP plant may reduce the biocoal production cost and increase operabilityby omitting the complex heat recovery scheme required for a standalone HTC plant.IGCC power plants with carbon capture are more efficient when employing entrainedflow gasification fired on torrefied wood or HTC biocoal than when using fluidized bedgasification of raw wood. However, the higher efficiency of the IGCC cannot compensatefor the conversion losses of the biomass upgrading. Moreover, the carbon capture ratefor scenarios with biomass upgrading is only 66–69%, compared to 82–86% for the directfluidized bed gasification of the raw biomass. The unit cost of electricity generated bythe BECCS plants is strongly dependent on the CO2 price. The results indicate that ifthe carbon price is sufficiently high to incentivize CCS from fossil fuels, then favourableBECCS configurations are also close to economic viability.

Contents

List of figures v

List of tables vii

Nomenclature xiii

1 Introduction 1

2 Background 52.1 Biomass as an energy source . . . . . . . . . . . . . . . . . . . . . . . . . . . 5

2.1.1 Properties of biomass as a fuel for combustion and gasification . . . 62.1.2 Biomass ressources . . . . . . . . . . . . . . . . . . . . . . . . . . . 7

2.2 Biomass upgrading technologies . . . . . . . . . . . . . . . . . . . . . . . . . 132.2.1 Pelletization . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 162.2.2 Torrefaction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 172.2.3 Hydrothermal carbonization . . . . . . . . . . . . . . . . . . . . . . 212.2.4 Anaerobic digestion . . . . . . . . . . . . . . . . . . . . . . . . . . . 35

2.3 Bioenergy with carbon capture (BECCS) . . . . . . . . . . . . . . . . . . . 402.3.1 CCS technologies . . . . . . . . . . . . . . . . . . . . . . . . . . . . 402.3.2 Gasification and pre-combustion carbon capture . . . . . . . . . . . 422.3.3 Technology options for BECCS . . . . . . . . . . . . . . . . . . . . 48

3 Scope, methodology and assumptions 493.1 Scenarios and simulation cases . . . . . . . . . . . . . . . . . . . . . . . . . 503.2 Supply chain cost and GHG emissions of biofuel production . . . . . . . . 54

3.2.1 Biomass cultivation and harvest . . . . . . . . . . . . . . . . . . . . 543.2.2 Storage . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 573.2.3 Transport . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 583.2.4 Processing at the biofuel plant . . . . . . . . . . . . . . . . . . . . . 603.2.5 Combustion of the biofuel at a power plant . . . . . . . . . . . . . . 60

3.3 Process simulation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 623.3.1 Property methods . . . . . . . . . . . . . . . . . . . . . . . . . . . . 623.3.2 Models for process units and general assumptions . . . . . . . . . . 64

3.4 Economic assessment . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 713.4.1 Estimation of the capital investment . . . . . . . . . . . . . . . . . . 723.4.2 Carrying charges . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 743.4.3 Feedstock, auxiliary energy and other consumables . . . . . . . . . . 743.4.4 Labour . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 753.4.5 Maintenance material . . . . . . . . . . . . . . . . . . . . . . . . . . 753.4.6 Carbon certificates . . . . . . . . . . . . . . . . . . . . . . . . . . . . 75

i

3.4.7 Levelized product cost . . . . . . . . . . . . . . . . . . . . . . . . . 763.5 Exergy and exergoeconomic analysis . . . . . . . . . . . . . . . . . . . . . . 76

3.5.1 Application of exergy-based analysis to biomass upgrading . . . . . 783.5.2 Application of exergy analysis to the BECCS plants . . . . . . . . . 80

3.6 Definitions for yields and efficiencies . . . . . . . . . . . . . . . . . . . . . . 80

4 Biomass upgrading processes 834.1 Drying . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 834.2 Pelletization . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 86

4.2.1 Design and simulation models of a wood pelletizing plant . . . . . . 864.2.2 Energy and carbon balance . . . . . . . . . . . . . . . . . . . . . . . 874.2.3 GHG emissions . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 884.2.4 Economic performance . . . . . . . . . . . . . . . . . . . . . . . . . 89

4.3 Torrefaction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 914.3.1 Design and simulation model of a torrefaction plant . . . . . . . . . 914.3.2 Energy balance, carbon balance and GHG emissions . . . . . . . . . 934.3.3 Exergy analysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 934.3.4 Economic performance . . . . . . . . . . . . . . . . . . . . . . . . . 94

4.4 Anaerobic digestion . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 974.4.1 Design and simulation model of anaerobic digestion with biomethane

production . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 974.4.2 Design and simulation model of anaerobic digestion with biofuel

pellet production . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 974.4.3 Simulation model of the anaerobic digestion reaction . . . . . . . . 994.4.4 Energy balance . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 994.4.5 Carbon balance . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1014.4.6 GHG emissions . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1014.4.7 Exergy analysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1034.4.8 Economic performance . . . . . . . . . . . . . . . . . . . . . . . . . . 104

4.5 Hydrothermal carbonization . . . . . . . . . . . . . . . . . . . . . . . . . . 1064.5.1 General considerations for HTC as a pretreatment for combustion

and gasification . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1064.5.2 Considerations for an industrial-scale plant design . . . . . . . . . . 1094.5.3 Simulation model of the HTC reaction . . . . . . . . . . . . . . . . 1104.5.4 Design and simulation model of the HTC base case . . . . . . . . . . 1124.5.5 Energy balance . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1144.5.6 Carbon balance . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1144.5.7 GHG emissions . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1154.5.8 Environmental considerations . . . . . . . . . . . . . . . . . . . . . 1164.5.9 Exergy analysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1184.5.10 Economic performance . . . . . . . . . . . . . . . . . . . . . . . . . . 1194.5.11 Exergoeconomic analysis . . . . . . . . . . . . . . . . . . . . . . . . . 1224.5.12 Parameter studies and alternative designs . . . . . . . . . . . . . . . 1254.5.13 HTC with anaerobic digestion of the wastewater . . . . . . . . . . . 1324.5.14 HTC integrated with CHP . . . . . . . . . . . . . . . . . . . . . . . 1364.5.15 HTC and subsequent combustion in a biomass fired CHP plant . . . 141

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4.6 Comparison of the biomass upgrading processes . . . . . . . . . . . . . . . 1444.6.1 Exergetic performance of the biofuel production . . . . . . . . . . . 1444.6.2 Conversion chain efficiency for the combustion of upgraded biofuels

in existing power stations . . . . . . . . . . . . . . . . . . . . . . . . 1454.6.3 Economic performance . . . . . . . . . . . . . . . . . . . . . . . . . 1474.6.4 GHG balance and mitigation cost . . . . . . . . . . . . . . . . . . . . 1514.6.5 Identifying the most suitable plant size . . . . . . . . . . . . . . . . . 1534.6.6 Solid biofuels versus biomethane . . . . . . . . . . . . . . . . . . . . 157

5 Bioenergy with carbon capture (BECCS) 1615.1 Syngas production . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 161

5.1.1 Design and simulation model of syngas production with entrainedflow gasifier . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 162

5.1.2 Design and simulation model of syngas production with fluidizedbed gasifier . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 163

5.1.3 Gasifier efficiency and syngas composition . . . . . . . . . . . . . . 1655.1.4 Energy balance . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1675.1.5 Carbon balance . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1695.1.6 Exergy analysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 170

5.2 IGCC power plants . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1725.2.1 Design and simulation model of an IGCC with entrained flow gasifier 1725.2.2 Design and simulation model of an IGCC with fluidized bed gasifier 1735.2.3 Energy and carbon balance . . . . . . . . . . . . . . . . . . . . . . . 1765.2.4 Exergy analysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1795.2.5 Economic performance . . . . . . . . . . . . . . . . . . . . . . . . . . 180

5.3 Comparison with other biomass, coal and gas to power technologies . . . . 183

6 Conclusions and outlook 1896.1 Comparison of different biomass upgrading technologies . . . . . . . . . . . 1896.2 Detailed analysis of the HTC plant design . . . . . . . . . . . . . . . . . . . 1926.3 Bioenergy with carbon capture and storage (BECCS) . . . . . . . . . . . . 1946.4 Outlook . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 196

A Modelling assumptions 199A.1 Supply chain of biofuel production . . . . . . . . . . . . . . . . . . . . . . . 199

A.1.1 Storage . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 199A.1.2 Transport distance and cost . . . . . . . . . . . . . . . . . . . . . . 199A.1.3 Supply chain GHG emissions . . . . . . . . . . . . . . . . . . . . . . 204

A.2 Process simulation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 204A.2.1 Modelling assumptions . . . . . . . . . . . . . . . . . . . . . . . . . . 204

A.3 Assumptions for the cost estimates . . . . . . . . . . . . . . . . . . . . . . . 218A.3.1 Adjustment to year 2010 € and plant location . . . . . . . . . . . . 218A.3.2 Equipment cost CBM . . . . . . . . . . . . . . . . . . . . . . . . . . 218A.3.3 Offsite costs . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 236A.3.4 Costs of auxiliary energy and consumables . . . . . . . . . . . . . . . 237A.3.5 Operating labour . . . . . . . . . . . . . . . . . . . . . . . . . . . . 237A.3.6 Levelized product costs . . . . . . . . . . . . . . . . . . . . . . . . . 239

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A.4 Exergetic efficiency definitions and auxiliary costing equations . . . . . . . 239A.4.1 Upgrading plants . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 239A.4.2 BECCS plants . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 239

B Biomass upgrading plant data 243B.1 Wood pelletizing plant models . . . . . . . . . . . . . . . . . . . . . . . . . 243

B.1.1 Cost data . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 243B.2 Torrefaction plant model . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 245

B.2.1 Simulation data from TOR-1.0 . . . . . . . . . . . . . . . . . . . . 245B.2.2 Cost data . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 246

B.3 Anaerobic digestion plant models . . . . . . . . . . . . . . . . . . . . . . . . 248B.3.1 Model for the separation of press fluid and press cake for ADP-3.0 . 248B.3.2 Digestion model . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 248B.3.3 Simulation data from ADM-3.0 and ADP-3.0 . . . . . . . . . . . . 251B.3.4 Digester design and heat loss . . . . . . . . . . . . . . . . . . . . . . 253B.3.5 Cost data . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 254

B.4 HTC plant models . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 257B.4.1 HTC reaction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 257B.4.2 Simulation data from HTC-3.00 . . . . . . . . . . . . . . . . . . . . 258B.4.3 Simulation data from HTC-5.00 . . . . . . . . . . . . . . . . . . . . 259B.4.4 The fate of nutrients from EFB in HTC-5.00 . . . . . . . . . . . . . 261B.4.5 Cost data from the HTC base design . . . . . . . . . . . . . . . . . . 262B.4.6 Exergy analysis and exergoeconomic analysis . . . . . . . . . . . . . 265B.4.7 Alternative flowsheet designs HTC-3.20 to HTC-3.80 . . . . . . . . 267B.4.8 Summarized results from HTC-3.01 to HTC-3.90 . . . . . . . . . . 270B.4.9 Simulation data from HTC-3.30 and HTC-3.60 . . . . . . . . . . . 272B.4.10 Cost data from HTC-3.30 and HTC-3.60 . . . . . . . . . . . . . . . 274B.4.11 HTC-3.90 with AD of the waste water . . . . . . . . . . . . . . . . 278B.4.12 Flowsheet designs of the integrated HTC and CHP plants CHPB-3.1

and CHPB-3.3 . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 278B.4.13 Simulation data from CHPB-3.2 . . . . . . . . . . . . . . . . . . . . 280B.4.14 Cost data from CHPB-3.1, CHPB-3.2 and CHPB-3.3 . . . . . . . . 281B.4.15 HTC followed by the subsequent combustion in a CHP plant . . . . 287

B.5 GHG emissions from biofuel production (WP, TOR, HTC, ADM, ADP) . . 287B.6 Combustion of upgraded biofuels in existing power stations . . . . . . . . . 288

C Data of the BECCS plants 291C.1 Syngas production . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 291C.2 IGCC . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 297

C.2.1 Simulation data from EF-IGCC-HTC-1 . . . . . . . . . . . . . . . . 297C.2.2 Simulation data from FB-IGCC-wood-1 . . . . . . . . . . . . . . . . 302C.2.3 Simulation data from FB-IGCC-wood-0 . . . . . . . . . . . . . . . . 307C.2.4 Cost data from EF-IGCC-HTC-1 . . . . . . . . . . . . . . . . . . . 310C.2.5 Cost data from FB-IGCC-wood-1 . . . . . . . . . . . . . . . . . . . . 314C.2.6 Investment cost summary for the IGCC plants . . . . . . . . . . . . 318

Bibliography 319

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List of Figures

2.1 Van Krevelen diagram of experimental data for HTC. . . . . . . . . . . . . 242.2 Relation between energy yield and HHV for HTC and torrefaction. . . . . . 282.3 Molar O/C ratio of biocoal from beech wood, cellulose and lignin carbonized

for 4 and 17 h at different temperatures. . . . . . . . . . . . . . . . . . . . . 292.4 Simplified flow diagram of gasification with pre-combustion carbon capture. 42

3.1 Conversion pathways in which biofuels replace fossil fuels in existing powerstations. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 53

3.2 Conversion pathways with BECCS and their associated respective referencescenarios with fossil fuel use and biomass without carbon capture. . . . . . 53

3.3 Schematic of the biofuel supply chain. . . . . . . . . . . . . . . . . . . . . . 553.4 Schematic for the efficiency definitions of biocoal utilization in a power plant. 613.5 Efficiency and turbine inlet temperature (TIT) of gas turbine systems in

relation to their capacity. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 70

4.1 Drying systems: high temperature drying, low temperature drying, SSD. . . 844.2 Flowsheet of the wood pellet plant with CHP, WP-1.2. . . . . . . . . . . . 874.3 Flowsheet of the torrefaction plant TOR-1.0. . . . . . . . . . . . . . . . . . 924.4 Flowsheet of anaerobic digestion with biomethane production ADM-3.0. . . 984.5 Flowsheet of anaerobic digestion with solid biofuel pellets ADP-3.0. . . . . 984.6 Effect of the degree of carbonization on the exergetic efficiency of combus-

tion and gasification. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1074.7 Relation between biomass moisture and energetic efficiency of a boiler with

exhaust gas temperatures of 160°C and 60°C, for the combustion of rawbiomass and for the conversion chain of HTC and subsequent combustion. . 108

4.8 Flowsheet of the HTC base case HTC-3.00. . . . . . . . . . . . . . . . . . . 1134.9 Production costs of biocoal from EFB, factoring in certificates for avoided

GHG emissions from EFB dumping. . . . . . . . . . . . . . . . . . . . . . . 1234.10 Cost of exergy destruction CD,k and cost associated with the equipment

items Zk for HTC-3.00-s and HTC-1.00-s. . . . . . . . . . . . . . . . . . . . 1244.11 Flowsheet of HTC-3.60. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1304.12 Flowsheet of HTC-3.70. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1314.13 Flowsheet of the combined HTC and CHP plant CHPB-3.2. . . . . . . . . 1374.14 Flowsheets of the biomass fired CHP plants. . . . . . . . . . . . . . . . . . . 1424.15 Exergy of biomass and electricity consumption, waste streams and main

sources of exergy destruction for selected biofuel plant models. . . . . . . . 1454.16 Levelized fuel plus carbon certificate costs in relation to the CO2 price. . . 1484.17 Biofuels from wood: revenues and production cost per unit of biofuel. . . . . 149

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4.18 Biofuels from waste and grass: revenues and production costs per unit ofbiofuel. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 150

4.19 GHG mitigation cost and GHG reduction per unit feedstock biomass. . . . 1534.20 Specific investment costs of biofuel production plants in relation to the plant

capacity. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1544.21 Break-down of biocoal cost versus plant capacity for HTC-3.00. . . . . . . 1554.22 Biofuel production cost and GHG mitigation cost for different feedstock

supply scenarios. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1564.23 Surplus cost per MWhel when using biofuels instead of fossil fuels and GHG

mitigation costs, both in relation to the feedstock cost. . . . . . . . . . . . . 159

5.1 Flowsheet of the syngas production process with entrained flow gasifier EF-HTC-1. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 162

5.2 Flowsheet of the syngas production process with fluidized bed gasifier FB-wood-1. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 164

5.3 Methane content of the syngas as a function of the gasification temperature.Experimental data and simulation data calculated with approach tempera-tures and under chemical equilibrium. . . . . . . . . . . . . . . . . . . . . . 167

5.4 Energy input for the production of 1 MJ of clean, de-carbonized syngas. . . 1685.5 Breakdown of the electricity consumption for the syngas production cases. 1695.6 Exergy flows for the syngas production cases. . . . . . . . . . . . . . . . . . 1715.7 Flowsheet of the IGCC with entrained flow gasifier EF-IGCC-HTC-1. . . . 1745.8 Flowsheet of the IGCC with fluidized bed gasifier FB-IGCC-wood-1. . . . . 1755.9 Composite curves for the HRSG cases. . . . . . . . . . . . . . . . . . . . . 1795.10 Exergy flows for selected IGCC cases. . . . . . . . . . . . . . . . . . . . . . 1805.11 Breakdown of the module cost (CBM) per unit of electricity production for

the IGCC cases. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1815.12 Efficiency and specific investment cost for IGCC plants analyzed in this

work and for comparable plants from other studies. . . . . . . . . . . . . . . 1835.13 Breakdown of the COE for various power plant types. . . . . . . . . . . . . 1855.14 COE in relation to the carbon price. . . . . . . . . . . . . . . . . . . . . . . 186

B.1 Flowsheet of HTC-5.00. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 259B.2 Flowsheet of HTC-3.20. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 268B.3 Flowsheet of HTC-3.30. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 268B.4 Flowsheet of HTC-3.40. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 269B.5 Flowsheet of HTC-3.50. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 269B.6 Flowsheet of HTC-3.80. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 269B.7 Flowsheet of CHPB-3.1. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 279B.8 Flowsheet of CHPB-3.3. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 279B.9 Flowsheet of SC-3.1.4. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 284

C.1 Flowsheet of the syngas production process with autothermal reformingFB-wood-3. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 291

C.2 Flowsheet of the IGCC with fluidized bed gasifier without carbon capture,FB-IGCC-wood-0. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 306

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List of Tables

2.1 Water and ash content of biomass and waste resources and bituminous coal 72.2 Waste biomass potentials in Germany . . . . . . . . . . . . . . . . . . . . . 102.3 Spatial and seasonal distribution of selected biomass resources. . . . . . . . 122.4 Selected waste streams, potentially suitable for HTC, in Berlin. . . . . . . . 122.5 Thermochemical conversion processes for lignocellulosic biomass . . . . . . . 152.6 Carbon yield of the three product phases from HTC. . . . . . . . . . . . . . 232.7 Mass yield of gaseous byproducts from HTC of poplar, old leaves and oil

palm empty fruit bunches at 220°C and 4 h. . . . . . . . . . . . . . . . . . . 252.8 Concentrations of identified substances in the process water from the HTC

of empty fruit bunches and poplar shavings at 220°C and 4 h. . . . . . . . . 262.9 Ash content of feedstock and biocoal for HTC from wood, empty fruit bun-

ches and digestate. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 262.10 Biogas yields and conversion efficiencies for anaerobic digestion using diffe-

rent feedstocks. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 38

3.1 Scenarios for biomass upgrading. . . . . . . . . . . . . . . . . . . . . . . . . 513.2 BECCS scenarios and their reference cases. . . . . . . . . . . . . . . . . . . 523.3 Biomass feedstocks used in this work. . . . . . . . . . . . . . . . . . . . . . 573.4 Transportation distances in the biomass supply chain. . . . . . . . . . . . . 583.5 Assumptions for the calculation of the transport distance to the upgrading

plant processing wood or grass, and the resulting required catchment area. 593.6 Cost of solid biofuel shipment by bulk cargo to a European harbour. . . . . 603.7 Assumptions on fossil fuel prices. . . . . . . . . . . . . . . . . . . . . . . . . 623.8 Steam cycle parameters. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 703.9 Key assumptions for the economic assessment. . . . . . . . . . . . . . . . . 723.10 Assumptions for the calculation of the total capital investment (TCI). . . . 74

4.1 Exergetic fuel and exergy destruction and exergy losses of different dryingsystems for wood chips and biocoal. . . . . . . . . . . . . . . . . . . . . . . 85

4.2 Energy demand and energetic efficiency for the four wood pelletizing cases. 884.3 Conversion chain efficiency from wood to electricity when the wood pellets

are co-combusted in a coal-fired power plant. . . . . . . . . . . . . . . . . . 894.4 Investment and levelized wood pellet production costs using short rotation

wood. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 904.5 Levelized wood pellet production costs using forest residues in North Ame-

rica and then shipping the pellets to Europe. . . . . . . . . . . . . . . . . . 904.6 Mass yields, HHV and energy yield for the torrefaction simulation model. . 924.7 Energy demand and energetic efficiency of the torrefaction plant. . . . . . . 934.8 Exergy balance for the torrefaction plant TOR-1.0. . . . . . . . . . . . . . 94

vii

4.9 Exergy losses and exergy destruction for the torrefaction plant TOR-1. . . 954.10 Investment costs for the torrefaction plants processing short rotation wood. 964.11 Plant capacities, specific investment costs and levelized production costs of

torrefied pellets from short rotation and from forest residues. . . . . . . . . 964.12 Key data from the anaerobic digestion model. . . . . . . . . . . . . . . . . 1004.13 Energy balance of the anaerobic digestion plant models. . . . . . . . . . . . 1004.14 Carbon balance of the anaerobic digestion plant models. . . . . . . . . . . 1014.15 GHG emissions from the anaerobic digestion plants. . . . . . . . . . . . . . 1024.16 Exergy balance of the anaerobic digestion plants models. . . . . . . . . . . . 1034.17 Exergy losses and exergy destruction for the anaerobic digestion plant models.1044.19 Specific investment and levelized biofuel production costs for the anaerobic

digestion plants. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1044.18 Investment costs for the anaerobic digestion plants. . . . . . . . . . . . . . 1054.20 Biocoal composition and yields in the simulation cases. . . . . . . . . . . . . 1124.21 Energy balance and efficiencies for the HTC base design for various feedstocks.1154.22 Carbon yields of biocoal and byproducts for the HTC base design. . . . . . 1154.23 GHG emissions from HTC biocoal production for selected cases. . . . . . . 1164.24 Exergy balance of HTC-3.00. . . . . . . . . . . . . . . . . . . . . . . . . . . 1184.25 Exergy losses and exergy destruction for HTC-3.00. . . . . . . . . . . . . . 1184.26 Investment cost for HTC-3.00-s and HTC-3.00-m. . . . . . . . . . . . . . . 1204.27 Plant capacities, specific investment and levelized biocoal production costs

for the HTC base design. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1214.28 Energetic efficiency, TCI, specific TCI per unit of biocoal and levelized

biocoal production costs for different operating parameters and flowsheetdesigns. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 126

4.29 Alternative designs for HTC plants. . . . . . . . . . . . . . . . . . . . . . . 1284.30 Anaerobic digestion of HTC waste water: simulation model and measured

data. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1344.31 Energetic fuels and products, important energy flows and energetic efficien-

cies of the integrated HTC and CHP systems. . . . . . . . . . . . . . . . . . 1384.32 Exergy analysis results for the integrated HTC and CHP systems. . . . . . 1394.33 Investment costs of the CHPB systems, and levelized costs ci,bc assigned to

the biocoal production. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1414.34 Simulation cases for HTC plants with subsequent combustion and for the

combustion of raw or dried biomass. . . . . . . . . . . . . . . . . . . . . . . 1424.35 Conversion chain efficiency for HTC with subsequent combustion and for

the combustion of raw or dried biomass. . . . . . . . . . . . . . . . . . . . . 1434.36 Efficiencies of wood pellets, biocoal, bituminous coal and biomethane com-

bustion in conventional power stations. . . . . . . . . . . . . . . . . . . . . 1464.37 GHG emissions related to biofuels and fossil fuels per unit of biofuel. . . . . 1524.38 Biomass and biocoal road transport costs for a 100 km distance. . . . . . . 155

5.1 Simulation cases for syngas production with CCS. . . . . . . . . . . . . . . 1625.2 Gasification agents, gasifier outlet composition, HHV of the raw gas, cold

gas efficiency and exergetic efficiency of the gasifiers. . . . . . . . . . . . . . 166

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5.3 Cold gas efficiencies for the production of clean, decarbonized syngas, forthe syngas production process and for the overall conversion chain includingthe pretreatment of the biomass. . . . . . . . . . . . . . . . . . . . . . . . . 168

5.4 Carbon balance for the syngas production cases including pretreatment,based on the carbon contained in the raw biomass. . . . . . . . . . . . . . . 170

5.5 Simulation cases for IGCC plants. . . . . . . . . . . . . . . . . . . . . . . . 1725.6 Efficiencies and carbon capture rates for the IGCC cases and their respective

conversion chains. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1765.7 Comparison of the modelling results for biomass and coal-fired IGCC from

this work with those from other studies. . . . . . . . . . . . . . . . . . . . . 1785.8 Levelized cost of electricity for the IGCC plants. . . . . . . . . . . . . . . . 1825.9 Key assumptions and results for the reference plants. . . . . . . . . . . . . 184

A.1 Storage dry matter loss and cost. . . . . . . . . . . . . . . . . . . . . . . . . 199A.2 Assumptions for road transport costs. . . . . . . . . . . . . . . . . . . . . . 200A.3 Costs for road transport. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 200A.4 Assumptions for the calculation of the transport distance of waste to the

upgrading plant. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 202A.5 Assumptions and results for the shipping cost. . . . . . . . . . . . . . . . . 203A.6 GHG emissions from transport. . . . . . . . . . . . . . . . . . . . . . . . . 204A.7 Parameters for the heat capacity correlation. . . . . . . . . . . . . . . . . . 205A.8 Efficiencies of turbomachinery. . . . . . . . . . . . . . . . . . . . . . . . . . 206A.9 Minimum temperature differences of heat exchangers. . . . . . . . . . . . . 207A.10 Overall heat transfer coefficients for heat exchangers. . . . . . . . . . . . . . 207A.11 Heat losses of process units. . . . . . . . . . . . . . . . . . . . . . . . . . . 207A.12 Assumptions for heat loss calculations of the HTC reactor. . . . . . . . . . 208A.13 Pressure losses of process units. . . . . . . . . . . . . . . . . . . . . . . . . 209A.14 Milling electricity demand. . . . . . . . . . . . . . . . . . . . . . . . . . . . 210A.16 Inert gas consumption for lock hoppers and piston feeders. . . . . . . . . . 211A.15 Electricity consumption of biomass pressurizing. . . . . . . . . . . . . . . . 212A.17 Modelling assumptions for mechanical dewatering. . . . . . . . . . . . . . . 213A.18 Modelling assumptions for driers. . . . . . . . . . . . . . . . . . . . . . . . 213A.19 Modelling assumption for combustion. . . . . . . . . . . . . . . . . . . . . . 214A.20 Modelling assumptions for gasification. . . . . . . . . . . . . . . . . . . . . 214A.21 Data from IGT/Renugas gasifier and resulting approach temperatures. . . 215A.22 Modelling assumptions for acid gas removal units. . . . . . . . . . . . . . . 215A.23 Modelling assumptions for gas cleaning and conditioning equipment. . . . 216A.24 Modelling assumptions for gas turbine systems, and resulting efficiencies for

natural gas-fuelled operation. . . . . . . . . . . . . . . . . . . . . . . . . . . 217A.25 Modelling assumptions for the air separation units. . . . . . . . . . . . . . 217A.26 Modelling assumptions for the reciprocating engine. . . . . . . . . . . . . . 217A.27 Cost data for heat exchangers. . . . . . . . . . . . . . . . . . . . . . . . . . 220A.28 Cost data for turbines, pumps, compressors, fans and motors. . . . . . . . . 221A.29 Cost data for boilers and furnaces. . . . . . . . . . . . . . . . . . . . . . . . 222A.30 Cost data for driers. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 223A.31 Cost data for mechanical dewatering equipment. . . . . . . . . . . . . . . . 224A.32 Cost data for screening and sizing equipment. . . . . . . . . . . . . . . . . 224

ix

A.33 Cost data for feeding and pressurization equipment. . . . . . . . . . . . . . 225A.34 Cost data for solids handling equipment at IGCC plants. . . . . . . . . . . 226A.35 Cost data for pelletizing and pellet handling equipment. . . . . . . . . . . 227A.36 Cost data for tanks. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 227A.37 Cost data for reactors. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 228A.38 Cost data for gasifiers. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 229A.39 Cost data for particulate removal and miscellaneous gas cleaning equipment. 230A.40 Cost data for desulphurization and acid gas removal equipment. . . . . . . 231A.41 Cost data for water gas shift and methane steam reforming reactors. . . . 232A.42 Cost data for gas turbine systems. . . . . . . . . . . . . . . . . . . . . . . . 233A.43 Cost data for miscellaneous power plant equipment. . . . . . . . . . . . . . 234A.44 Cost data for auxiliary units: air separation, waste water treatment, engine,

refrigeration. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 235A.45 Cost data for IGCC offsite cost. . . . . . . . . . . . . . . . . . . . . . . . . 235A.46 Scaling exponents α for offsite costs, and offsite cost as % of PEC in the

reference cases of the upgrading plants. . . . . . . . . . . . . . . . . . . . . 236A.47 Reference cases for the calculation of the offsite cost. . . . . . . . . . . . . 236A.48 Cost of auxiliary energy and other consumables. . . . . . . . . . . . . . . . 237A.49 Hourly cost of labour and scaling exponents for labour requirement. . . . . 238A.50 Requirement of plant operators per shift. . . . . . . . . . . . . . . . . . . . 238A.51 Requirement of workers for biomass yard, product handling, feedstock sizing

and sorting, per shift. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 238A.52 Escalation rates. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 239A.53 Exergetic efficiency definitions for torrefaction and anaerobic digestion equip-

ment. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 239A.54 Exergetic efficiency definitions and auxiliary costing equations for HTC-1.00

and HTC-3.00. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 241

B.1 Investment costs for wood pellet plant models using SR wood. . . . . . . . 243B.2 WP-1.0-s equipment list with investment costs. . . . . . . . . . . . . . . . 244B.3 WP-1.2-m equipment list with investment costs. . . . . . . . . . . . . . . . 244B.4 Composition (d.b.) of the torrefied wood for the torrefaction simulation

model. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 245B.5 TOR-1.0 flow stream data. . . . . . . . . . . . . . . . . . . . . . . . . . . . 245B.6 TOR-1.0-m equipment list with investment costs. . . . . . . . . . . . . . . 247B.7 Mass flows into and composition of press fluid and press cake, for simulation

model ADP-3.0, and based on data from literature. . . . . . . . . . . . . . . 249B.8 Data for the anaerobic digestion model for simulation cases ADP-3.0, ADM-

3.0 and ADM-3.1, and based on data from literature. . . . . . . . . . . . . 250B.9 ADM-3.0 flow stream data. . . . . . . . . . . . . . . . . . . . . . . . . . . 251B.10 ADP-3.0 flow stream data. . . . . . . . . . . . . . . . . . . . . . . . . . . . 252B.11 Assumptions for the digester design. . . . . . . . . . . . . . . . . . . . . . . 253B.12 Assumptions for the digester heat loss calculation. . . . . . . . . . . . . . . 253B.13 Digester design and heat loss calculation. . . . . . . . . . . . . . . . . . . . 253B.14 ADM-3.0-s equipment list with investment costs. . . . . . . . . . . . . . . 255B.15 ADP-3.0-s equipment list with investment costs. . . . . . . . . . . . . . . . 256B.16 Mass yields and biocoal composition (d.b.) from the HTC reactor model. . 257

x

B.17 Composition of the unspecified dissolved organic compounds from the HTCreactor model. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 257

B.18 HTC-3.00 flow stream data. . . . . . . . . . . . . . . . . . . . . . . . . . . 258B.19 Composition (d.b.) and heating value of fibres and shells used as boiler fuel. 259B.20 HTC-5.00 flow stream data. . . . . . . . . . . . . . . . . . . . . . . . . . . 260B.21 Nutrients in EFB, HTC waste water, and GHG emissions of the equivalent

mineral fertilizers. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 261B.22 Monetary value of the nutrients in the EFB and in the HTC waste water,

calculated with fertilizer prices for Malaysia and Indonesia. . . . . . . . . . 261B.23 Investment cost from the HTC base cases. . . . . . . . . . . . . . . . . . . . 262B.24 HTC-3.00-s equipment list with investment costs. . . . . . . . . . . . . . . 263B.25 Component results from the exergoeconomic analysis of HTC-3.00-s. . . . . 265B.26 Flow stream results from the exergoeconomic analysis of HTC-3.00-s. . . . 266B.27 Component results from the exergoeconomic analysis of HTC-1.00-s. . . . . 267B.28 Results from HTC-3.20 to HTC-3.90. . . . . . . . . . . . . . . . . . . . . . 270B.29 Results from HTC-3.01 to HTC-3.13. . . . . . . . . . . . . . . . . . . . . . 271B.30 HTC-3.30 flow stream data. . . . . . . . . . . . . . . . . . . . . . . . . . . . 272B.31 HTC-3.60 flow stream data. . . . . . . . . . . . . . . . . . . . . . . . . . . . 273B.32 HTC-3.30-s equipment list with investment costs. . . . . . . . . . . . . . . 274B.33 HTC-3.60-s equipment list with investment costs. . . . . . . . . . . . . . . 276B.34 Flow stream data for the anaerobic digestion in HTC-3.90. . . . . . . . . . 278B.35 CHPB-3.2 flow stream data. . . . . . . . . . . . . . . . . . . . . . . . . . . 280B.36 Investment cost summary of the integrated HTC and CHP systems. . . . . 281B.37 Annual cost flows for the standalone CHP plant and the integrated plants. 282B.38 Equipment list with investment costs of the standalone CHP plant. . . . . 283B.39 Energy balance of HTC followed by the subsequent combustion in a CHP

plant and direct combustion of raw biomass, SC-1.0.1 to SC-3.1.4. . . . . . 285B.40 Exergy balance of HTC followed by the subsequent combustion in a CHP

plant and direct combustion of raw biomass, SC-1.0.1 to SC-3.1.4. . . . . . 286B.41 GHG emissions for biofuel production for WP, TOR, HTC, ADM and ADP. 287B.42 Water content, density, calorific value, energy density and costs for 100 km

road transport for different types of raw biomass and upgraded biofuels. . 288B.43 Key results from selected biomass upgrading scenarios. . . . . . . . . . . . 289

C.1 Energy balance for the syngas production. . . . . . . . . . . . . . . . . . . . 292C.2 Exergy balance for the syngas production. . . . . . . . . . . . . . . . . . . . 293C.3 EF-HTC-1 flow stream data. . . . . . . . . . . . . . . . . . . . . . . . . . . 294C.4 FB-wood-1 flow stream data. . . . . . . . . . . . . . . . . . . . . . . . . . . 295C.5 FB-wood-3 flow stream data. . . . . . . . . . . . . . . . . . . . . . . . . . . 296C.6 EF-IGCC-HTC-1 flow stream data from the gasification section. . . . . . . 297C.7 EF-IGCC-HTC-1 flow stream data from the air separation, acid gas re-

moval and gas turbine system. . . . . . . . . . . . . . . . . . . . . . . . . . 298C.8 EF-IGCC-HTC-1 flow stream data from the steam cycle. . . . . . . . . . . 299C.9 EF-IGCC-HTC-1 mechanical and electrical work data. . . . . . . . . . . . 300C.10 EF-IGCC-HTC-1 steam production and consumption. . . . . . . . . . . . 301C.11 FB-IGCC-wood-1 flow stream data from the gasification section, air separ-

ation, acid gas removal and gas turbine system. . . . . . . . . . . . . . . . 302

xi

C.12 FB-IGCC-wood-1 flow stream data from the steam cycle. . . . . . . . . . . 303C.13 FB-IGCC-wood-1 mechanical and electrical work data. . . . . . . . . . . . 304C.14 FB-IGCC-wood-1 steam production and consumption. . . . . . . . . . . . 305C.15 FB-IGCC-wood-0 flow stream data from the gasification section, air separ-

ation, acid gas removal and gas turbine system. . . . . . . . . . . . . . . . 307C.16 FB-IGCC-wood-0 flow stream data from the steam cycle. . . . . . . . . . . 308C.17 FB-IGCC-wood-0 mechanical and electrical work data. . . . . . . . . . . . 309C.18 EF-IGCC-HTC-1 equipment list with investment costs. . . . . . . . . . . . 310C.19 EF-IGCC-HTC-1 heat exchanger list with investment cost. . . . . . . . . . 312C.20 FB-IGCC-wood-1 equipment list with investment costs. . . . . . . . . . . . 314C.21 FB-IGCC-wood-1 heat exchanger list with investment cost. . . . . . . . . . 317C.22 Investment cost summary for IGCC plants. . . . . . . . . . . . . . . . . . . 318

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Nomenclature

Roman symbolsA area (heat exchanger surface in [m2], biomass harvest area in [ha])a yearb specific GHG emissions [kgCO2,eq/MJ]C cost flow [€/h]c mass-specific heat capacity [kJ/kg/K]c specific cost per unit of mass, energy or exergy (units as stated)CD cost rate of exergy destruction [€/h]cGHG GHG mitigation cost [€/tCO2,eq]d diameter [m]E exergy flow [MW]e specific exergy per unit of mass [kJ/kg]ED rate of exergy destruction [MW]EL rate of exergy loss [MW]fd overdesign factorfHTC degree of carbonizationfk exergoeconomic factor of component kfM material factorfp pressure factorfT temperature factorh height [m]i interest rate per annumm mass flow [kg/s], [kg/h]n number (of equipment items or plant operators)p pressure [bar]psat saturation pressure [bar]rk relative cost difference between the costs per unit of fuel and product

exergy of component kT temperature [°C]t time [h], [s]T0 ambient temperature [°C]T d dew point temperature [°C]

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T sat saturation temperature [°C]V volume flow [m3/h]V volume [m3]W electrical or shaft work [kW]w water content (mass fraction of fresh matter)x arithmetic meanX capacity of an equipment item (units as stated)xoDM organic dry matter content [kg/kg]yk exergy destruction ratio of component kZ cost rate associated with investment and maintenance of an equipment

item [€/h]

Greek symbolsα scaling exponent for adjusting equipment cost or labour requirement to

the plant capacityΔp pressure difference, pressure loss [bar]ε exergetic efficiencyη energetic efficiencyγc carbon yield (ratio of carbon in upgraded biofuel to carbon in feedstock)γe energy yield (ratio of upgraded biofuel energy to feedstock energy, based

on HHV unless otherwise stated)γm mass yield (ratio of upgraded biofuel mass to feedstock mass)ρ mass density [kg/m3]

Subscripts0 at ambient conditions (15°C, 1.013 bar, unless otherwise stated)auxfuel auxiliary fuelCC conversion chainCO2,eq CO2 equivalents according to the GWP100 warming potentialCOD chemical oxygen demanddaf dry and ash-freeD (exergy) destructiond design (capacity of plant equipment)DM dry mattere energy (based on HHV unless otherwise stated)el electricalev evaporated waterex exergyF fuel (exergy)

xiv

FM fresh matter (including moisture)HHV higher heating valuelg liquid and gas phaseLHV lower heating valuem massoDM organic dry matterPE primary energyP product (exergy)PP power plantref referencesim simulation (capacity of plant equipment in simulation)s solid (biofuel, biomass)STP at standard temperature and pressure (0 °C, 1.013 bar)th thermaltot total (overall plant or system)unit per equipment unit

SuperscriptsCH chemical (exergy)PH physical (exergy)

AbbreviationsAD anaerobic digestionADM anaerobic digestion with biogas upgrading to biomethaneADP anaerobic digestion of press juice with pellet production from press cakeAFUDC allowances for funds used during constructionAGR acid gas removalASU air separation unitBECCS bioenergy with carbon capture and storageBSP biomass steam processing (torrefaction in superheated steam)CBM module costCCGT combined cycle gas turbine power plantCCS carbon capture and storageCELF constant escalation levelization factorCGE cold gas efficiencyCHPB combined heat, power and biocoal productionCHP combined heat and power (cogeneration)COD chemical oxygen demandCOE cost of electricity

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CRF capital recovery factorCS carbon steeldaf dry and ash-freed.b. on dry basisDM dry matterEFB empty fruit bunches (from palm oil production)EF entrained flow (gasification)ETS emissions trading systemFB fluidized bed (gasification)FCI fixed capital investmentFFB fresh fruit bunches (from palm oil production)FM fresh matter (including moisture)FR forest residuesFTS Fischer-Tropsch synthesisFWPH feedwater preheaterGHG greenhouse gasesGWP100 100-year global warming potentialHHV higher heating valueHMF hydroxymethylfurfuralHP high pressureHPLC high performance liquid chromatographyHRSG heat recovery steam generatorHRT hydraulic retention timeHTC hydrothermal carbonizationHT high temperatureHW hot waterHX heat exchangerIGCC integrated gasification combined cycle (power plant)LHV lower heating valueLP low pressureLTD low temperature dryingLT low temperatureMOW (source-separated) municipal organic wasteMP medium pressureNPV net present valueoDM organic dry matterOLR organic loading rateO&M operation and maintenance

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PC pulverized coal-fired power plantPEC purchased equipment costPE primary energyPGW park and gardening wastePGW-70 park and gardening waste with a water content of 70%POME palm oil mill effluentRM Malaysian ringgit (currency)SR short rotationSSD superheated steam dryingSS stainless steelTCI total capital investmentTOC total organic carbon (dissolved in HTC waste water)TOMres residual total organic matter (dissolved in HTC waste water), without

acetic and formic acidsTOM total organic matter (dissolved in HTC waste water)TOR torrefactionTRR total revenue requirementUSD US dollarVOC volatile organic compoundsVP very low pressurew.b. on wet basisWP wood pelletizingwt% % by weight

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1 Introduction

The use of biomass in the European Union has been increasing in the last decades and,according to the climate and energy policy goals of the EU and its member states, shallcontinue to grow [1]. For example, the German Biomass Action Plan aims at expandingthe bioenergy use from 792 PJ in 2007 to 1309 PJ in 2020, corresponding to 11% of theGerman primary energy demand [2].

While EU policy in the last ten years has focussed in part on the production of transportfuels from biomass [3, 4], the greenhouse gas (GHG) mitigation potential of these fuels iscontroversial. Several recent studies conclude that electricity and heat from biomass offera greater GHG reduction than do transport biofuels [5–7]. This is partly due to the factthat heat and power production allows the use of perennial crops and wastes as feedstocks,whereas current technologies for transport fuel production require more environmentallyharmful crops like rapeseed and maize. Increasing the use of waste biomass is consideredthe most important option to expand the bioenergy supply in an economically and envir-onmentally viable way [2, 5]. Perennial crops such as short rotation coppice may provideadditional potential [2]. However, a report by the German Advisory Council on the En-vironment concludes that the potential of waste and energy crops produced in Germanyis limited to 5% of the primary energy demand, thus biomass imports are inevitable if thepolicy goals are to be met [6]. For the EU, a deficiency in wood supply of around 870 PJcan be expected for 2020 due to the increasing demand for bioenergy [8] (based on datafrom [9]).

Given the relative scarcity in Europe and large potentials in other world regions such asNorth America and Russia, biomass is increasingly becoming an internationally tradedcommodity fuel [1, 10]. In that context, processed biomass including wood pellets isgaining in importance. Due to its higher energy density, improved handling and betterstorability, pretreated biomass is much better suited for long-distance transport than rawbiomass [10].

In terms of biomass conversion technologies, co-combustion with coal has been identifiedas a low cost short term option to employ bioenergy [11]. What is more, it provides fuelflexibility for existing power stations. According to the IEA World Energy Outlook 2011,80% of the CO2 emissions allowed until 2035 under a 450 ppm greenhouse gas stabilizationtarget are already “locked-in” by existing power plants and other facilities [12]. Using theexisting infrastructure with carbon neutral biomass instead of fossil fuel could thereforebe crucial to achieving GHG reduction targets.

The same properties which make pelletized biomass favourable for international trade alsomake it well-suited as a power plant fuel. In the last decade, several European electricitycompanies have started to co-fire wood pellets in their power station and are planning toincrease their pellets use [13].

1

Chapter 1 Introduction

Instead of simple pelletizing, the properties of biomass can be further improved by chemicalpretreatment such as torrefaction and hydrothermal carbonization (HTC). By destroyingthe fibrous structure of the biomass and removing oxygen-containing functional groups,these processes increase the carbon content and calorific value, and make the biomass waterrepellent and easier to mill. This more coal-like upgraded biomass may be processed in theexisting coal handling infrastructure and feeding systems and is therefore attractive forco-firing. While torrefaction, a form of mild pyrolysis, requires prior drying, HTC takesplace in water and is therefore well suited for wet feedstocks, including various forms ofwaste biomass.In the near term, upgrading technologies can help to unlock new biomass potentials byenabling the efficient use of wet waste biomass, improving biomass properties for interna-tional trade, and facilitating their co-firing in existing power stations. In the long term,they may play an important role for bioenergy with carbon capture and storage (BECCS).When the feedstock biomass is grown sustainably, BECCS leads to a net removal of CO2from the atmosphere. Several recent studies have found that such negative emissionstechnologies may be crucial to avoid dangerous climate change [14–17]. In order to limitthe global temperature rise to 2°C above pre-industrial levels, atmospheric greenhouse gas(GHG) concentration must be stabilized at a level of 400–450 ppm CO2 equivalents orpossibly lower, depending on the climate sensitivity. GHG targets below 400 ppm requireglobal CO2 emissions to become close to zero or even negative. Several global energymodels show that ambitious GHG stabilization targets of below 400 ppm are unlikely tobe achieved without BECCS, and that targets around 450 ppm are cheaper to meet whenBECCS is employed [16, 17]. When CO2 capture rates are sufficiently high, BECCS canbecome the prevailing biomass conversion technology [17, 18].CCS technology is likely to be first developed for coal, and may not be easily transferableto biomass. Vergragt et al. argue that the investment in capture plants and CO2 transportand storage infrastructure may lead to a reinforcement of fossil fuel lock-in of the energysystem [19]. Upgrading technologies like torrefaction and HTC, which produce a coal-likebiofuel and thereby facilitate the use of coal-based CCS infrastructure in combination withrenewable resources, may alleviate the risk of fossil fuel lock-in.Hydrothermal carbonization has recently attracted considerable interest in Germany, dueto its capacity to process most types of biomass and convert it into a high quality solidbiofuel. While a substantial amount of laboratory-scale experimental research on thechemical reaction itself has been published, assessments of the potential technical andeconomic performance of industrial-scale HTC plants are still lacking in the scientificliterature. One aim of this work is to fill this gap.Some key research questions are therefore:

• Does HTC make sense as a biomass pre-processing step for combustion and gasific-ation from a thermodynamic and economic point of view?

• Which technological niches can HTC occupy in the current energy system? Whichare the best feedstocks? What is a suitable plant capacity?

• How does HTC compare to other, more established, upgrading technologies?• Can HTC and/or other biomass upgrading technologies improve the technical and

economic performance of BECCS schemes?

2

Introduction

As of early 2013, there are no commercial HTC plants in operation, thus the efficiencyand cost of the overall process are unknown. An essential element of this work thereforeis the development of a flowsheet design for an industrial-scale HTC plant, which allowsto estimate the auxiliary energy consumption, energetic efficiency and investment cost ofthe overall process. The simulation package Aspen Plus is used to model the HTC plant,other upgrading technologies for comparison and the BECCS plants.

The next chapter provides background on the availability of the biomass resources con-sidered in this work, on biomass upgrading processes and on BECCS. The upgrading tech-nologies analyzed in this work, besides HTC, comprise wood pelletizing, torrefaction, andanaerobic digestion. Integrated gasification combined cycle (IGCC) with pre-combustioncarbon capture is explored as a BECCS technology.

In chapter 3, the employed methods and modelling assumptions are explained. The modelsused for the process units such as reactors, heat exchangers, compressors etc. in the AspenPlus simulation are described (section 3.3) and the cost functions and assumptions forthe investment cost estimates and the calculation of the levelized product costs are given(section 3.4). Exergy analysis and exergoeconomic analysis are applied as design tools todiscover potentials for improvement in the HTC plant design. These exergy-based methodsand their application to the evaluation of the biomass upgrading processes are explainedin section 3.5.

In order to evaluate a suitable capacity range for HTC plants, the transport cost of thebiomass has to be taken into account. An increasing plant capacity on the one hand leadsto lower specific investment and labour costs due to economy-of-scale effects, but on theother hand requires a larger catchment area and thus longer supply lines. The model usedto estimate the transport cost and GHG emissions as a function of the plant capacity isdescribed in section 3.2.

The results for the biomass upgrading plants are presented in chapter 4. The processesare discussed in terms of energy and exergy balance, and specific GHG emissions andcost per unit of biofuel are given. For HTC, several variations in flowsheet configurationand operating parameters are modelled in order to seek the best plant design and assessthe sensitivity of efficiency and costs (section 4.5.12). Integration of HTC with anaerobicdigestion for waste water treatment, and with a biomass-fired combined heat-and-power(CHP) plant for heat recovery are explored in sections 4.5.13 and 4.5.14, respectively. Insection 4.5.15, the conversion chain efficiency of HTC followed by combustion is comparedto the direct combustion of raw biomass. The chapter closes with a comparison of thedifferent upgrading technologies, focussing on the economic viability and GHG mitigationcosts when substituting fossil fuels in existing power stations.

In chapter 5, the results for the BECCS plants are discussed. Energy, exergy and carbonbalances are first analyzed for several plant designs and simulation cases of the syngasproduction from HTC biocoal, torrefied wood and raw wood (section 5.1). Selected syngasproduction cases are then extended to a model of an overall IGCC plant, for which costestimates are made and the levelized cost of electricity is calculated (section 5.2). At theend of the chapter, the modelled BECCS plants are compared to other types of powerplant with and without carbon capture in respect of overall carbon capture rate and costs.

Chapter 6 offers a conclusion and indicates the scope for future work.

3

Chapter 1 Introduction

4

2 Background

This chapter first discusses the properties of different biomass feedstocks and their implic-ations for combustion and gasification. This serves to illustrate the need for upgradingtechnologies. Further details are provided on the availibility of the feedstocks consideredin this work. The current state of research and deployment is reviewed for the analyzedbiomass upgrading and BECCS technologies.

In addition, publications by the author related to this work comprise [20–23].

2.1 Biomass as an energy source

Biomass and waste currently cover around 10% of the global primary energy demand.With its contribution of 60% to the total renewable energy, biomass can be considered asubstantial renewable energy source [5, page 33]. However, more than 80% is currentlybeing used in traditional cooking and heating applications with a very low efficiency [5].Within the EU-27, the contribution of biomass to the total primary energy amounted to4.6% in 2006 [24, page 11].

The extraction of useful energy from biomass is versatile regarding feedstock, product andconversion technology. It includes, for instance, such different processes as combustion ofwood in domestic heating applications, anaerobic digestion of sewage sludge for heat andpower generation, and fermentation of maize to produce bioethanol. Because differenttypes of biomass have different properties, the conversion technologies have to be tailoredto the feedstock. In 2007, the dominating bioenergy technology in Germany was biodieselproduction (257 PJ) from domestic rapeseed and imported oil seeds including soy and palmoil. This was followed by domestic heating applications (233 PJ) and power, district heator CHP plants (approximately 174 PJ) fuelled on wood chips [25]. Biogas to electricityusing maize and, to a lesser extend, grass silage, contributed 77 PJ in 2007 [25] yet hasbeen rapidly expanding in the last few years. In 2011, biogas contributed 67% of theelectricity generation from biomass in Germany [26].

The GHG mitigation potential varies widely depending on the selected feedstock and con-version technology. Sterner and Fritsche analyzed 74 bioenergy pathways for Germanyand found that GHG emissions range from plus 257 kg/GJ to minus 190 kg/GJ CO2equivalents compared to fossil reference systems [7]. Careful analysis of all available feed-stocks and conversion pathways is therefore essential in order to focus policy support onpathways with a high GHG reduction.

5

Chapter 2 Background

2.1.1 Properties of biomass as a fuel for combustion and gasification

First generation biofuels require feedstocks rich in oil (for biodiesel) or sugar and starch(for bioethanol). Oil, sugar and starch, however, are only present in large quantities in thefruits, seeds and tubers of certain plants. Wood, leaves, stalks and haulms predominantlyconsist of the polymers cellulose (50%), hemicellulose (25%) and lignin (20%) [27]. Woodyand herbaceous biomass is therefore also referred to as lignocellulosic biomass.

Cellulose is formed from linear chains of up to 15␣000 glucose molecules. It is insolublein water, relatively resistant against enzymatic decomposition, and responsible for thefibrous nature of biomass. Hemicelluloses are branched heteropolymers of different sugars.They possess side chains which make them partly water soluble, and are important forthe moisture absorption capacity of plants. Lignin forms a complex highly linked 3-dimensional network structure of aromatic components and is responsible for the rigidityof wood. With 64%, it has the highest carbon content of the biomass polymers [27, 28]. Inconversion processes operated at moderate temperatures of around 200°C, such as HTCand torrefaction, the various polymers can behave quite differently. This is discussed inthe sections on the respective technologies.

In comparison to coal, biomass has a much higher O/C ratio. For bituminous coal, thecarbon mass fraction of the dry and ash-free matter (daf) is around 80–90% and theoxygen fraction 2–10% [29]. For biomass, carbon contributes only 44–63% and oxygen40–47% [28, pages 343, 360]. Because of the high O/C ratio, the higher heating valueof biomass only reaches 18.5–21 MJ/kg (daf) [28, page 360], while for coal, it is 33–37MJ/kg [29]. The biomass volatile matter content of 70–85% [28, page 360] is much higherthan that of bituminous coal (35%) and lignite (50%) [30], which makes biomass morereactive. Biomass is also prone to biological decomposition, which facilitates enzymaticconversion processes such as anaerobic digestion and fermentation, but also poses problemsfor storage.

One of the biggest disadvantages for biomass as a fuel is its high moisture content. Mostbiomass except straw and certain energy grasses like miscanthus consists of more than50% water. Part of this water is bound chemically or locked in capillaries and cell walls[28, page 315].

The ash content of biomass is usually lower than that of coal, except from waste feedstockslike sewage sludge, source-separated municipal organic waste (MOW) and roadside grass.The ash and water contents of different biomass and waste resources are summarized inTable 2.1. The sulphur content of biomass at typically 0.02–0.2% is low. The sulphur,nitrogen, potassium and chlorine contents are generally higher for fertilized agriculturalcrops than for unfertilized woody biomass [28]. This makes wood more straightforwardto use as a fuel than herbaceous biomass, which tends to cause problems with corrosion,fouling and emissions [31].

The ash melting temperature is dependent on the ash composition, especially the po-tassium to calcium (K/Ca) ratio. While the ash melting temperature of wood is between1200–1400°C, in the same range as bituminous coal, that of straw is around 900°C [30].Empty fruit bunches from palm oil production, an extensive waste biomass resource, havean ash sintering temperature as low as 620°C [32]. A low ash melting temperature lim-its the maximum operating temperature of fluidized bed combustion and gasification and

6


gives rise to the problem of slag deposits [28].

Apart from causing harm in combustion facilities, the minerals contained in the biomassash, especially phosphorus, potassium and magnesium, are valuable plant nutrients [28].Returning the biomass ash to the forests or agricultural land from where the biomass wasextracted will close these mineral cycles. This is not possible if biomass is co-combustedwith coal [31]. An assessment of the environmental performance of biomass (co-)firingrequires that the effects on the mineral cycles be considered.

Table 2.1: Water and ash content of biomass and waste resources and bituminous coal[28, 30, 33–37].

water (w.b.) ash (d.b.)wood, fresh 45–65% 0.5–2%grass, fresh 65–80% 5-6%roadside grass 45–75% 18–25%grass, ensilage 60–85%miscanthus (perennial grass) 15–45% 1.5–4%straw 10–20% 5–12%haulm of sugar-beet and potato 75–80% 4–5%dead leaves 54–60% 20–36%source-separated municipal organic waste (MOW) 70% 20–40%sewage sludge, dewatered 55–70% 39–53%empty fruit bunches from palm oil production 65% 6%bituminous coal 2–15% 5–16%

Because of the high moisture content, the combustion efficiency of biomass is generally low.Moreover, fuel preparation processes like milling and feeding and the supply logistics arealso affected by the chemical and mechanical properties of the biomass. The inhomogen-eous and fibrous structure poses problems for feeding systems [11, 38], and the electricitydemand for milling the fibrous materials to small particle sizes is high [39]. Biologicaldegradation during storage can cause multiple problems including significant dry matterloss, health risks from fungi and self ignition [28].

The low calorific value and often low bulk density make biomass transport over longdistance uneconomic. Electricity generation from biomass therefore occurs mostly in small-scale plants with relatively low efficiencies. The electrical efficiency is 20–30% (based onLHV) for typical biomass-fired power plants with capacities of 20–50 MWel, and 10–15%for anaerobic digestion followed by a reciprocating engine [40]. One widely discussed optionto use biomass more efficiently is co-combustion in coal-fired power stations [11, 38, 41, 42].Generally, the maximum co-firing rate is limited to about 20% for wood, 5% for strawand 4% for sewage sludge for technical reasons [43]. Pretreatment of the biomass withpelletization, torrefaction or HTC should increase the potential co-firing ratio.

2.1.2 Biomass ressources

Biomass ressources for energy conversion processes comprise on the one hand energy crops,i.e. plants that are explicitly cultivated for the purpose of producing electricity, heat or

7


transport fuels, and on the other hand waste biomass from industrial processes, agricultureand households.

The technical potential for energy crops depends on the availability of arable land, and istherefore linked to the land area required for food production, which in turn is dependenton population levels, nutrition habits and land productivity. Estimates of the globalbiomass potential vary widely. A literature survey of 17 studies found a range from 47EJ/a to 450 EJ/a for the year 2050 [44].

Common energy crops such as maize and rapeseed are highly controversial, due to poten-tial land use competition with food production, negative effects on biodiversity and highGHG emissions related to the high input agriculture systems required for their cultiva-tion. A high nitrogen fertilizer demand is especially problematic, because of the energyconsumption embodied in the fertilizer production and its environmental impacts includingemissions of the greenhouse gas N2O, nitrate leaching into groundwater and eutrophicationof water-bodies [28, page 112]. Land use change is a complex issue with great import-ance for the GHG balance and the general environmental performance of energy crops.Meyer-Aurich et al. report that electricity production from biogas with maize as the solefeedstock may result in higher GHG emissions than using natural gas, when former grass-land is converted to agricultural land to grow the crop [45]. Conversely, the cultivation oflow-input perennial grassland species on abandoned degraded agricultural land has beenshown to lead to a net sequestration of carbon due to root mass accumulation [46]. Theadverse effects of land use change are especially detrimental when virgin forests or bogsare converted into agricultural land, as is the case for palm oil production in Indonesia [5].

Perennial woody and grassy crops, such as short rotation poplar or willow, switchgrassand miscanthus, are generally less environmentally damaging due to their low fertilizer de-mand, positive effects on the soil carbon balance and erosion mitigation [5, 28]. Moreover,especially when they are grown on inferior land, the risk of competition with food produc-tion is limited [5]. Due to their high lignocellulose content, however, they require differentconversion processes than oil and starch containing plants.

Waste biomass is generally considered to be the most environmentally benign bioenergyresource. The global potential is estimated to be in the range of 40–170 EJ/a [40].

Due to their better environmental performance, only perennial energy crops and wasteare considered in this work. Based on a literature search, wood, grass, municipal organicwaste and palm oil residues are selected as potential feedstocks for HTC and other up-grading technologies. In the following, some detail on these types of biomass is provided.In section 2.1.2.5, their seasonal and spacial distributions are discussed, as these have asignificant influence on the supply chain logistics and the sizing of the processing facilities.

2.1.2.1 Wood

Wood is the most commonly used bioenergy resource for combustion and gasification. Itcan be provided as forest residues, waste wood from industry and households, and as adedicated energy crop. The sum of forest residues and waste wood available for bioenergyin Germany amounts to approximately 29 million tonnes dry matter, or 574 PJ, per year[47, page 16]. These traditional potentials are already exploited to a great extent, thus

8


an increasing wood demand has to be partly covered by imports and/or the cultivation ofwoody crops.

Import from overseas is often conducted in the form of pellets rather than raw woodchips, due to their better transport properties. Major exporting regions are Russia, witha potential of around 20 million tonnes of forest residues per year and a limited domesticconsumption, and America, responsible for 46% of the worldwide wood chips production[48].

Short rotation cropping is still in its infancy in Germany and most other European coun-tries [49], but attracts an increasing interest. The plantation area in Germany increasedfrom 1200 ha in 2008 to approximately 4000 ha in 2012 [50]. In 2009, a German electricitycompany planted 238 ha of paulownia in South Spain, one of the biggest short rotationplantations worldwide, in order to secure the feedstock supply for a nearby biomass-firedpower station [51]. Fast growing species like willow and poplar with a harvest intervalof 3–10 years are the favoured cropping systems [28, pages 88–91]. Nitrogen fertilizer isgenerally not required [28, page 90].

Compared to other types of biomass, wood is well suited for combustion and gasifica-tion, due to its relatively low moisture, very low ash content, and relative suitability fortransport and storage.

2.1.2.2 Grass

Permanent grassland is an environmentally beneficial form of land use because of its highsoil carbon storage, low agricultural inputs and high biodiversity [46, 52]. It is traditionallyused for cattle grazing, but due to the increasing use of arable crops for livestock feed,grassland areas are gradually being abandoned. Grassland which is not harvested becomesovergrown with woods and shrubs [53]. It is estimated that 13–22% of the permanentgrassland in the EU, 9–15 million hectares, will not be required anymore for livestockfeeding in 2020 [54]. Because of its ecological value, preservation of grassland areas is apolitical goal in Germany and other European countries. New utilization concepts suchas harvesting the grass as bioenergy resource are therefore desired [52].

Grass is mostly discussed as a feedstock for biogas production. Utilization as a combustionfuel is also possible but requires prior drying. One problem for biogas production is thatbiodiversity conservation requires a late cut due to bird breeding times, but advancinglignification leads to a low digestibility and subsequent low biogas yield when the harvestis delayed [55, 56].

Storage generally is conducted as ensilage, which further increases the water content.

Aside from grass cultivated as an energy crop, grass also accrues as waste from roadsideedges and from municipal parks and sport fields [57].

2.1.2.3 Waste biomass in Germany

Waste biomass comprises forest residues, agricultural residues such as straw and liquidmanure, residues from food processing and pulp and paper production, and municipal

9


waste streams including sewage sludge, household waste and park and gardening wastes.The total amount of waste biomass is Germany is not statistically recorded [58]. Estimatesof the exploitable potential, with consideration of technical and ecological restrictions, arein the range of 60–80 million tonnes dry matter per year [33, 47, 58].

Since 2005, the landfill of untreated waste is prohibited in Germany [59]. Therefore,pathways for the utilization or disposal have been established for most types of waste.New technologies such as HTC compete with the current utilization or treatment processesand will only prosper if they are more cost efficient or explicitly incentivized, for examplebecause of a superior environmental performance.

Table 2.2 gives an overview of the major organic waste streams in Germany and theircurrent use. Only forest residues are currently deployed on a larger scale for energy.Most waste streams with a high moisture content are composted. Hay from landscapingand agricultural residues such as straw often remain on the field because recovery is tooexpensive [58]. Since composting suffers from high costs and difficulties in selling thecompost due to market saturation [33, pages 2, 138], alternative treatment technologiesare desirable.

Schuchardt and Vorlop assess the potential of waste biomass in Germany which could serveas a feedstock for HTC. They conclude that, because of the utilization schemes alreadyin place, the potentially available biomass (DM) for HTC is limited to an annual 1.0 Mtsewage sludge, 2.4 Mt organic waste and 1.6 Mt hay from conservation and landscaping[58].

Many types of waste have an ash content well above 20%, which impacts their suitabilityfor combustion and gasification.

Table 2.2: Waste biomass potentials in Germany based on [28, 33, 47, 58, 60, 61].

Mt/a DM current useforest residues 29 combustion for heat and power productionstraw 3.6–16 stable bedding, fertilizer, remains on fieldhaulm of sugar beet and potato 6.2 fodder production, remains on fieldliquid and solid manure 19.8 organic fertilizer, biogas productionsource-separated organic household waste 1.5–2.4 composting, biogas productionmunicipal park and gardening wastes 1.6 compostinghay from conservation areas 1.6 mulchingsewage sludge 2.2–3.5 combustion, biogas productionfood industry 0.6–1.4 fodder production, chemical industry

2.1.2.4 Palm oil residues

While due to the legislation most waste in Germany is utilized or adequately treated,dumping of organic residues is common practice in many countries. Converting thesewaste resources into useful energy would not only reduce GHG emissions by displacingfossil fuels, but also by preventing methane emissions from the rotting process. One suchexample are the empty fruit bunches (EFB) from palm oil production.

10


In 2009, worldwide palm oil production amounted to 45 million tonnes crude palm oil,with Indonesia (44 %) and Malaysia (43 %) being the main producers. The generationof one tonne of crude palm oil requires approximately 5 t of fresh fruit bunches (FFB),and produces around 3 t of liquid palm oil mill effluent (POME), 1.1 t of EFB, 0.6 t ofmesocarp fibres, 0.3 t of shells, and 0.4 t of kernel [62]. The kernel is usually sold as aresource for the production of palm kernel oil [62]. The energy content of the remainingsolid residues, namely EFB, fibres and shells, amounts to approximately 22.4 MJHHV pertonne of crude palm oil.

The energy demand of the palm oil process is covered by a steam turbine plant. Theboiler, fuelled on fibres and shells, produces steam at 17–25 bar, which is expanded to3–4 bar in a back-pressure steam turbine [32, 63]. The expanded steam is used for thesterilization of the FFB. The electricity demand amounts to approximately 20 kWh pertonne of FFB [62].

The boiler fuel demand amounts to around 8.9 MJHHV per tonne of crude palm oil andcan be covered by part of the fibres and shells. Excess fibres and shells are commonlysold [62], but EFB and POME impose a waste problem. POME is often is stored in openponds, and EFB are either used for mulching at the oil palm plantation or simply dumped.EFB dumps and POME ponds release large amounts of methane.

Using the EFB for mulching at the plantation has beneficial environmental effects includingthe return of nutrients to the land, improving the soil structure, increasing the moistureretention capacity, stimulating root growth, and decreasing leaching and soil erosion onsteep land [64]. However, it also attracts harmful beetles and snakes and a fungi whichcauses stem rot in the oil palm trees [37, 65]. High transportation and distribution costmay also frustrate the use of EFB for mulching [65].

Alternative applications for the EFB are therefore sought. The most straightforward op-tion of burning the EFB in the boiler for additional electricity generation is often notfeasible because many palm oil mills are not grid-connected and cannot export surpluselectricity [62, 66]. Besides, EFB are poorly suited as a boiler fuel because of their highwater and potassium content and their extremely low ash sintering temperature of ap-proximately 620°C [32]. Composting, biogas generation and the production of combustiblebriquettes are discussed in literature as potential treatment technologies [32, 37, 67, 68].

2.1.2.5 Spatial and seasonal distribution of biomass resources

The seasonal storage of biomass and its transport over long distances are costly due toits low energy density and its susceptability to biological degradation. The spatial andseasonal distribution of the available resources are therefore crucial for the feedstock supplylogistics and the siting and sizing of the conversion facilities.

Agricultural residues and energy crops are generally subject to strong seasonal fluctuationsdependent on the harvest times. Compared to residues, the per hectare yield of energycrops is generally higher.

Average yields per hectare for different types of biomass are reported in Table 2.3.

Waste streams from biomass processing industries offer the advantage that large amountsof waste are available as a point source, often year-round or at least several months per

11


year. Processing capacities of palm oil mills range from 20–90 t/h of FFB [62].1 With anannual capacity factor of 34–51% [66, 69–71]2 and a yield of 0.23 t EFB per tonne FFB[62], this results in 13–93 kt/a of EFB. The relatively low capacity factor is due to thefact that the palm oil mills run only 15–20 hours per day and at 70–75% capacity. FFBare available year round, with a seasonal fluctuation of ±26% [69, pages 13].

Table 2.3: Spatial and seasonal distribution of selected biomass resources. Values inbrackets indicate the potential for energy conversion, taking into account agriculturalrequirements such as humus regeneration and stable bedding [28, 57, 72–74].

yield per hectare[tDM/ha]

short rotation woody crops 6–18grass 2–12forest residues 1.3–3.3straw 5–7 (0.5–3)sugar beet haulm 10–12 (1–3)MOW, urban area with 2000 inhabitants per km2 0.6 1)

1) estimate based on an average generation of 94 kgFM/a source separated MOW plus park and gardeningwastes per person, and a water content of 70%

Waste streams from urban areas seem promising as a feedstock for HTC because thecollection infrastructure already exists. Relatively large amounts, including householdwaste, sewage sludge and food waste from groceries and restaurants, are available yearround. In rural areas, green waste is often composted by households [61] and may thereforebe difficult to mobilize for bioenergy conversion.

Table 2.4 summarizes the key waste streams in Berlin. Dead leaves fallen in autumn,which are collected for aesthetic reasons and road safety, offer a relatively large potential.However, their seasonality is problematic.

Table 2.4: Selected waste streams, potentially suitable for HTC, in Berlin [75].

waste stream ktFM/asource-separated organic household waste 52.8dead leaves 96.3park and gardening wastes 43.6fruit and vegetable residues 11.1food wastes 38.0sewage sludge (DM) 87.4

Assessments of the economic viability of biomass conversion plants which take into accountthe logistics of biomass supply have been presented in literature. For instance, Leible etal. [47] analyze the economic performance of fast pyrolysis and gasification plants insouth Germany using forest residues and straw. Wirth et al. [76] investigate potentialsites for HTC plants in the German federal state of Brandenburg using a mathematical

1Based on data from 12 Malaysian mills.2Based on data from several palm oil mills.

12

2.2 Biomass upgrading technologies

programming model based on the warehouse location problem. Straw, wood chips fromshort rotation forestry and forest residues are considered as feedstocks. Mager [77] analyzesthe potential of municipal waste usable for HTC in German cities. Hamelinck [78] assessesvarious energy conversion pathways using forest residues and energy crops from severalEuropean and overseas locations, and considers transport by train, truck and ship.


A variety of processes have been suggested to convert raw biomass into combustion fuels.One significant process design issue is whether the conversion reactions take place underdry conditions or in water. Processes taking place in a dry atmosphere generally requireprior drying of the feedstock. Since many biomass feedstocks have a high moisture content,the energy needed for the evaporation of the moisture can be detrimental to the efficiencyof the overall process. Processes taking place in water (referred to as hydrothermal)do not require evaporation of the feedstock moisture and are therefore well suited forwet feedstocks. Hydrothermal processes can be conducted in subcritical or supercriticalwater. In order to keep the water in a liquid state at the required reaction temperature, thepressure must be at or above the saturation pressure. This is one of the main draw-backsof hydrothermal processes, since feeding biomass against high pressures in a continuousprocess presents an engineering challenge.

Biomass conversion in dry atmosphere involves mostly pyrolysis and, if an oxidant ispresent, gasification reactions. In hydrothermal reactions, hydrolysis plays an importantrole in the decomposition of the biomass.

The role of water in hydrothermal reactions is complex. Water may act as a solvent,catalyst, reactant and as a medium for material and energy transport [79, 80]. Theproperties of water at above 200°C are very different from those of water at ambientconditions, including a lower dielectric constant, fewer and weaker hydrogen bonds, ahigher isothermal compressibility, and a dissociation constant 3 orders of magnitude higherthan that of water at ambient temperature [79, 80]. This leads to a completely differentbehaviour as a solvent: organic compounds become increasingly soluble with temperature,while the solubility of inorganic salts decreases [79]. At 250–350°C, the solvent propertiesof water come close to those of organic solvents at room temperature [80]. Reactionkinetics are also influenced by the presence of water. Unlike gas-gas or gas-solid reactions,reactions in a liquid are often diffusion-controlled, whereby the viscosity of the liquidmedium determines the rate of reaction [79]. When exothermal reactions take place inhydrothermal processes, the surrounding water prevents local temperature peaks [81],making the temperature control of the process somewhat easier.

Gaseous, liquid and solid energy carriers can all be produced via hydrothermal as wellas dry conversion pathways. Generally, higher temperatures favour gas yields and lowertemperatures favour solid yields. To generate equivalent products, hydrothermal processesrequire lower temperatures than processes in dry atmosphere. This is mostly due to thefact that decomposition by hydrolysis starts at a lower temperature than decomposition bypyrolysis. Yan et al. conducted fibre analysis of loblolly pine before and after hydrothermaltreatment and torrefaction and report that 100% of the hemicellulose and 42% of the

13


cellulose were converted in hydrothermal treatment at 230°C. During torrefaction at 250°C,only 35% of the hemicellulose and 21% of the cellulose were converted [82]. Biologicalprocesses such as anaerobic digestion, where the hydrolysis is catalyzed by enzymes, takeplace at temperatures below 100°C, but require very much longer residence times.

Although most biomass conversion processes are aimed at one particular product andattempt to optimize its yield and quality, they often produce considerable amounts ofbyproducts. For example, pyrolysis focussed on oil production also generates solid charand combustible gases, while hydrothermal carbonization focussed on char productionyields organic compounds dissolved in water and gaseous byproducts. Byproducts whichhave a low water content can usually be combusted to provide energy for the process.Byproducts dissolved in water or with a very high moisture content, such as the dissolvedorganic components in hydrothermal carbonization or digestate from anaerobic digestion,are more problematic. On the other hand, the fact that problematic inorganic compoundslike chlorine and potassium mostly remain in the liquid phase benefits the solid productof hydrothermal carbonization. Simply washing biomass in water followed by mechanicaldewatering to reduce the alkali content can be considered a treatment technology in itsown right for herbaceous biomass [83]. Washing straw at 50–60°C has been shown toremove most of the chlorine and potassium [84].

Some biomass conversion processes can be designed to generate particular target chemicalsrather than biofuels. Superheated steam pyrolysis, for example, has been shown to yieldphenol, furfural, guaiacol and glycolaldehyde [85]. Simple drying and pelletizing can alsobe viewed as an upgrading process, although it does not involve chemical reactions.

Table 2.5 (with no claim to completeness) gives an overview of conversion processes applic-able to lignocellulosic biomass. Catalysts or a modified pH value may allow for differentoperating conditions or different products than those given in the table. Anaerobic diges-tion and pelletization are the only two mentioned processes which are fully commercialized,the remaining processes are still under research and development.

The processes analyzed in this work, namely hydrothermal carbonization, torrefaction,anaerobic digestion and wood pelletizing, are described in the following sections.

14


Tabl

e2.

5:T

herm

oche

mic

alco

nver

sion

proc

esse

sfo

rlig

noce

llulo

sic

biom

ass

[28,

86–1

02].

nam

ere

acti

onty

pem

ain

prod

uct

tem

pera

ture

pres

sure

resi

denc

eti

me

[°C

][b

ar]

torr

efac

tion

dry

solid

200–

300

~115

–60

min

flash

carb

oniz

atio

ndr

yso

lid~4

0010

–20

<30

min

stea

mex

plos

ion

stea

mat

mos

pher

eso

lid~2

2010

–30

(ste

am)

~10

min

biom

ass

stea

mpr

oces

sing

stea

mat

mos

pher

eso

lid25

0–35

0~1

30–1

50m

inhy

drot

herm

alca

rbon

izat

ion

hydr

othe

rmal

solid

180–

250

10–4

01–

12h

flash

pyro

lysi

sdr

yliq

uid

450–

500

~1~1

sfa

stpy

roly

sis

dry

oil/

char

slur

ry(b

iocr

ude)

~500

~1~1

shy

drot

herm

alliq

uefa

ctio

nhy

drot

herm

alliq

uid

250–

350

100–

200

15–6

0m

inga

sific

atio

ndr

yga

s70

0–16

001–

400.

5s–

hour

shy

drot

herm

alga

sific

atio

nhy

drot

herm

alga

s(C

H4

orH

2)

400–

700

200–

350

0.3–

4m

inan

aero

bic

dige

stio

nhy

drot

herm

al,

biol

ogic

alga

s(C

H4)

38–5

7~1

15–1

20d

15


2.2.1 Pelletization

The most straightforward way to convert biomass into a homogeneous fuel with a highenergy density is by simply drying and pelletizing it. Pelletization is the only technologyfor producing upgraded solid biofuels which is fully commercialized. It can therefore beregarded as the benchmark technology, with which all other technologies for producingupgraded biofuels have to compete.

The global wood pellet consumption in 2009 amounted to more than 12 Mt, double thatof 2006. Europe was the main consumer with 10.4 Mt [48]. There are separate marketsfor domestic and industrial wood pellets. Domestic wood pellets, used in small-scaleresidential heating devices, have higher quality standards defined by DIN EN 14961-2 andare 40–60% more expensive than industrial pellets.3 While domestic pellets are mostlyproduced regionally, industrial pellets are increasingly becoming an internationally tradedcommodity fuel [48]. In 2008, 19% of the US pellet production was exported to Europe[48]. DIN EN 14961-2 specifies that industrial wood pellets must have a lower heatingvalue of 16 MJ/kg or more, a water content less than 12% and an ash content below 3%.

Industrial pellets are used for co-firing in coal-fired power plants and district heatingfacilities. Around 50% of the world wood pellet production is combusted in big powerstations [13]. Co-firing initiatives by several large utilities in the EU indicate a probablefuture increase in the demand for industrial wood pellets [48]. For example, the powercompany Vattenfall aims to replace more than 50% of their hard coal consumption withbiomass by 2020 [103]. However, co-firing wood pellets is currently pursued mostly incountries where it qualifies for feed-in premiums or tradable renewable energy certificates,namely the Netherlands, Belgium and the UK [48]. The future of wood pellets co-firingtherefore depends strongly on carbon prices and other policy incentives.

Sawdust, a byproduct from lumber and pulp manufacturing, is currently the main feed-stock for wood pellet production. Due to an increased wood pellet demand and reductionsin lumber and pulp production in the last years [104], there is an increasing effort todecouple pellet production from the availability of sawmill residues [48]. Several largepellet plants processing chipped roundwood with a production capacity of 300–400 kt/ahave been installed in the USA in 2009 [10], the largest facility currently has a productioncapacity of 750 kt/a [105]. Typical plant capacities in Europe are considerably lower. Thefeedstock cost is the most important factor for the pellet production cost [106]. Due tothe unfolding shortage of wood residues in Europe, pelletization of other types of biomass,such as agricultural crops, straw and rape press cake, is being investigated [8, 107].

A high ratio of wood pellets co-firing in fluidized-bed and pulverized-fuel combustionsystems designed for coal is technically feasible. Existing coal-fired power stations haveeven been successfully operated on 100% wood pellets in demonstration projects [108].Generally, a biomass share of up to 10% has little or no influence on combustion behaviourand efficiency [108]. A high share of biomass, however, leads to a significant loss in capacity,and, since the plant is then operated in part-load, also in efficiency. For example, thecapacity of Tilbury power station, UK, was reduced by 30% when converted to 100%wood pellets, leading to an efficiency loss of 1.7 percentage points [109]. The need for

3In 2009, industrial pellet prices were around 140 €/t and prices for domestic pellets, bagged for retail,were 200–220 €/t [10].

16


modifications in the fuel storage and processing systems may also hinder the use of woodpellets in existing power stations [108]. Several big electricity companies in Germany aretherefore considering the use of upgraded pellets from torrefaction and HTC, which havea higher calorific value than wood pellets and facilitate the use of existing coal storageand processing infrastructure [110, 111].

Thanks to their favourable flow properties, pellets can be handled with equipment de-veloped for grain handling [112]. Due to their low moisture content they are less prone tobiological decay than raw biomass [28]. However, wood pellets need to be stored indoorsbecause contact with water and even absorption of moisture from the ambient air cancause them to swell and disintegrate [31].

The production of wood pellets comprises the unit operations of drying, comminution,conditioning, pelletizing, cooling and sieving [28]. Conditioning is the addition of steamif the moisture content after drying is too low to make the particles glue together [28].The increase in energy density relative to the raw material is mostly achieved by drying.Drying is also the most energy consuming step of pellet production if the feedstock hasa high moisture content, such are raw wood. The thermal energy for the driers usuallyderives from the combustion of bark.

2.2.2 Torrefaction

Torrefaction is a form of mild pyrolysis, conducted at a relatively low temperature levelof 200–300°C and at atmospheric pressure. The residence time is typically 15–60 minutes.During torrefaction, volatile compounds are removed from the feedstock, leading to adecrease in the O/C ratio. The product, torrefied wood (sometimes referred to as biocoal4),therefore has a higher calorific value than the raw biomass.

Research on torrefaction dates back to the 1930s in France, but little was published aboutthat early work [113]. One demonstration plant with a production capacity of 12 kt/aof torrefied wood was built and operated in the 1980s by the French company Pechiney.Forest residues served as the feedstock, and the product was mostly used as a coke sub-stitute in the metallurgical industry. The plant was decommissioned after a few years foreconomic reasons [114].

In the last few years, interest in torrefaction has been revived. A main contributor hasbeen the Energy Research Centre of the Netherlands (ECN), who have been working onthe topic since 2002 [113] and have published numerous articles and reports [113–123].

Bench-scale torrefaction experiments have mostly focussed on different types of wood[88, 122, 124–129] and herbaceous biomass with a low moisture content such as straw,reed canary grass and miscanthus [122, 130–132]. Some types of biomass waste have alsobeen torrefied, namely coconut shell [126], roadside grass [121], bagasse [122, 128], wastesfrom the palm oil industry [133, 134], hydrolysis residues from ethanol production [135],chicken manure and sewage sludge [136].

As of early 2012, there are no commercial-scale torrefaction plants in operation. ECNoperates a small pilot plant with a capacity of 100 kg/h [122]. Several companies are de-

4In this work, the term biocoal is used for the product of HTC, while the product of torrefaction isreferred to as torrefied wood.

17


veloping torrefaction and are operating pilot plants with a capacity of 1–5 t/h. Overviewsof these technology firms and their current projects can be found in [87, 137, 138]. Thefirst industrial-scale torrefaction plants with a capacities of 20–60 kt/a are currently underconstruction [139, 140].

2.2.2.1 Reaction chemistry

The torrefaction process can be subdivided into two steps: the depolymerization of thebiomass macromolecules and the subsequent charring reactions of the intermediates [125].The depolymerization includes dehydration, decarboxylation and deacetylation reactions[141]. Light volatiles (mono- and polysaccharides, dehydrosugars) are formed in the pro-cess [88]. At typical torrefaction temperatures, hemicellulose is the main reactant [115].During the decomposition of the hemicellulose, hydrophilic oxygen-containing groups (hy-droxyl, carbonyl and carboxyl) are removed and replaced by furan-aromatic and aliphaticstructures [122]. Thermal cleavage of carbon bonds of carboxylic groups leads to theformation of acids, which in turn catalyze dehydration and thermal cleavage reactions[88]. The degradation of intermediates catalyzed by mineral matter also plays a role [88].This makes herbaceous biomass with its higher ash content more reactive than wood [122].

The decomposition of hemicellulose takes place in the temperature range of 225–325°C,with hemicellulose from deciduous wood being more reactive than that of coniferous wood[141]. Decomposition of cellulose requires temperatures of 305–375°C, which is outside thetypical operating temperature range of torrefaction. Lignin decomposes slowly over a widetemperature range between 250–500°C [125]. Above 300°C, fast thermal cracking of cellu-lose may lead to tar formation, so an operating temperature below 300°C is recommended[125].

Torrefaction comprises both endothermic and exothermic reactions, and becomes increas-ingly exothermic with rising temperature.

2.2.2.2 Products

Depending on the operating conditions and the feedstock, 65–95% of the dry biomass re-mains in the torrefied wood, while the remaining 5–35% is converted to gaseous byproducts,referred to as torrefaction gas. The higher heating value of the torrefied wood is typic-ally 1–3 MJ/kg (d.b.) higher than that of the feedstock biomass, and the volatile mattercontent is decreased from 80% to 60–75% [141]. Approximately 75–95% of the feedstockcarbon and the feedstock energy content (HHV) remain in the torrefied wood.

The torrefaction gas consists mostly of water vapour, CO2, CO, acetic acid and methanol.Other compounds detected in smaller amounts include CH4, furfural, hydroxyacetone,formic acid, lactic acid, formaldehyde, acetaldehyde, acetone and phenol [114, 130, 141].Mass yields based on the dry biomass are 5–8% water, 1–4% CO2, and 1–7% acetic acid[114, 141]. The composition of the torrefaction gas varies with the feedstock. The gaseousproducts from straw contain more CO than those from wood. This is caused by thehigher mineral content (ash), which favours catalyzed reactions of CO2 and H2O withchar, yielding CO [88]. The composition of the organic byproducts also varies between

18


coniferous and deciduous wood, due to the different composition of their hemicelluloses[88].

The inorganic compounds (ash) remain in the solid phase. Since the organic mass decreasesduring torrefaction, the mass fraction of ash is higher in the torrefied biomass than in thefeedstock.

2.2.2.3 Influence of operating conditions

With increasing temperature and residence time, the mass yield of torrefied wood decreaseswhile the carbon content and calorific value increase. Since part of the feedstock carbonand hydrogen is consumed in the production of water, CO2 and other byproducts, a moresevere torrefaction results in a lower energy yield.

For industrial applications, the torrefaction must be severe enough to guarantee the desiredproduct properties, namely hybrophobicity and the destruction of the fibres. On the otherhand, excessive torrefaction should be avoided in order to obtain a high energy yield.Verhoeff et al. suggest torrefaction temperatures of 260°C for herbaceous biomass, 280°Cfor deciduous wood and 290°C for coniferous wood [122]. Arias recommends 240°C and aresidence time of 30 min for woody biomass [124].

2.2.2.4 Properties of torrefied wood

The decomposition of the hemicellulose has two beneficial effects in terms of fuel quality.Firstly, the depolymerization decreases the length of fibres and leads to the loss of thetenacious structure [122], with positive effects on grindability and rheological propertiessuch as flowability [129]. Secondly, the removal of oxygen-containing groups, which areresponsible for the high moisture absorption capacity of raw biomass, leads to hydro-phobic behaviour in the torrefied wood [122]. When exposed to water, torrefied woodpellets showed little swelling and no disintegration over the test period of 15 hours, whileconventional wood pellets swelled and disintegrated within minutes [39]. This is benefi-cial for storage, since torrefied pellets, unlike wood pellets, can be stored outside. Rezaet al., however, report that pellets from torrefied wood disintegrated within 5 minuteswhen immersed in water [142]. This different behaviour may be due to differences in thetorrefaction conditions and/or the pelletization process.

Compared to untreated wood, torrefaction reduces the electricity consumption for millingto 200 μm by 80–90% and increases the throughput of the mill by a factor of 10 [114].Simpler types of milling equipment such as cutting mills or jaw crushers can be employedinstead of the more costly hammer mills. Alternatively, size reduction can be achievedduring densification [114].

Optical microscopy shows that wood particles become more spherical during torrefaction.Fibres forming links between particles, as observed with raw wood, do not exist in thetorrefied material [124]. The loss of the fibrous nature and consequential improvementof the flow properties is especially important for pressurized entrained flow gasification,where finely milled fuel needs to be fluidized and pneumatically conveyed to the gasifier.Fluidization experiments show that, unlike raw wood, torrefied biomass can be fluidized,

19


although not as well as coal [115]. Torrefaction at higher temperatures may further improvethe fluidization behaviour [115].

To increase the energy density and thereby reduce transport and storage costs, Bergmansuggests to pelletize the torrefied wood [117]. The pelletization behaviour and pellet qual-ity (strength, hydrophobicity, grindability) are dependent on the torrefaction conditions[122]. A too severe torrefaction impairs the pelletization behaviour. This may be due todegradation of the lignin [39], which acts as a binding agent. Reza et al. report a veryhigh proneness to abrasion for pellets from wood torrefied at 300°C for 60 minutes, whichwould not meet various pellet standards [142]. Optimal torrefaction conditions whichsimultaneously achieve good pelletization behaviour, high pellet quality, good fluidizationbehaviour after milling and a high energetic efficiency of the torrefaction process are thesubject of further research.

2.2.2.5 Overall process design

An industrial torrefaction plant would typically comprise the unit operations drying andsizing of the feedstock biomass, torrefaction, cooling, sizing and densification of the torre-fied wood, and combustion of the gaseous byproducts from torrefaction.

According to manufacturer information, the optimal particle size for torrefaction is 25 x25 x 12 mm [143], but may depend on reactor type. The thickness must be below 40 mm[119].

Although the torrefaction reaction is slightly exothermic, thermal energy has to be sup-plied to the reactor to evaporate the remaining moisture and reach the required reactiontemperature. This can be achieved by indirect heat transfer, for example with thermal oil,or with direct heat transfer by recirculated torrefaction gas or superheated steam. Pro-posed reactor designs are often derived from drying or pyrolysis, and include the multiplehearth furnace, rotary drum reactor, screw conveyor reactor, and oscillating belt reactor[121]. Most studies on the simulation and design of industrial-scale plants are authoredby the Energy Research Centre of the Netherlands, who focus on a directly heated movingbed reactor [114, 117, 122, 144]. The majority of the remaining technology developersalso appear to employ directly heated reactors of various designs using hot gas as the heatsource. Torrefaction in superheated steam5 [89, 145] is another option for directly heatedreactors but seems to remain at an early state of research. Concepts pursuing indirectheat transfer include a heated rotating drum reactor [146].

Since the torrefaction gas contains CO and combustible organic compounds, it can beburned to provide thermal energy for the torrefaction reactor and/or drier. However, dueto its high water content it has a very low calorific value. When completely dry biomass istorrefied, the water content of the torrefaction gas is over 50% by weight [114]. In practice,the water content is significantly higher due to the moisture remaining in the biomass afterdrying. Trattner reports that a furnace for the co-combustion of biomass and torrefactiongas has been developed [146].

The capacity of a commercial plant is expected to be at least 50–100 kt/a of torrefiedproduct [121, 146].

5Also referred to as biomass steam processing (BSP).

20


2.2.3 Hydrothermal carbonization

Hydrothermal carbonization (HTC) — also referred to as hydrothermal pretreatment, wettorrefaction, artificial coalification and hot compressed water treatment — can be regardedas the wet equivalent of torrefaction. The solid product, biocoal, is similar to lignite. Thereaction takes place in water at temperatures of 180–250°C at or above saturation pressure.

HTC was first suggested by Friedrich Bergius in 1913 [147] as an experimental techniqueto mimick coal genesis under laboratory conditions. Much of the early work on HTC[147–153] was conducted in the field of coal petrology and focussed on understanding thereaction mechanisms which were thought similar to those that took place in the naturalformation of fossil coals. Many of the experiments conducted in that context employedrather high temperatures (up to 340°C) and residence times of up to several weeks, becausethey were aimed at reaching a degree of carbonization similar to that of bituminous coal.

Early technical applications of hydrothermal treatment include several pilot-scale andcommercial peat upgrading plants, in service between 1904 and 1964, and with capacitiesof 6 000 to 50 000 tonnes of product per year [154]. The peat upgrading process tookadvantage of the effect that the hydrothermal treatment destroys the colloidal structureof the peat, thereby releasing the chemisorbed water and the water enclosed in cells andenabling mechanical dewatering.

A renaissance in HTC research, centred in Germany, was triggered in 2006 by the sugges-tion of Markus Antonietti to apply HTC to biomass as a contribution to climate changemitigation [155]. To this end, HTC biocoal can be utilized in two ways. First, biocoal,unlike biomass, can provide a chemically stable carbon sink. This option is widely dis-cussed for biocoal from pyrolysis (biochar) [156] but has also been suggested for biocoalfrom HTC [157]. Second, biocoal can replace fossil coal in coal-fired power plants and thusavoid the use of fossil fuels and the related CO2 emissions.

In the last few years, HTC of biomass has been investigated in several research projects. Avariety of feedstocks have been successfully carbonized in laboratory scale experiments, in-cluding wood, straw, cut grass, dead leaves, municipal organic waste, fermentation residuesfrom anaerobic digestion, empty fruit bunches from palm oil production and distillersgrains [23, 158–165]. While these feedstocks result in a lignite like product, microalgaedisplay a very different reaction chemistry under hydrothermal conditions [166]. Carbon-ized microalgae are reported to resemble bituminous coal [166] or heavy fuel oil [167], withhigh levels of fatty acids absorbed onto the char [168].

Several companies, many of them located in Germany6, are trying to commercialize theHTC technology [170–176]. The first demonstration plant with a processing capacity of8400 t/a began operation in 2010 [175, 177]. As of late 2012, there are no commercial-scaleHTC plants in operation.

6A comprehensive list of HTC technology developers can be found at the website of the BundesverbandHydrothermale Carbonisierung (German Association for Hydrothermal Carbonization) [169]. As ofApril 2013, 18 member organizations are listed.

21



Hydrothermal carbonization is not a single reaction but a complex network of reactions,which has yet to be understood in detail. Reaction mechanisms include decarboxyla-tion, dehydration, condensation and aromatization [81]. Based on HTC experiments withmodel compounds such as glucose, xylose, cellulose and lignin, some understanding hasbeen developed on the reaction pathways of the individual polymers which constitute realbiomass [27, 162, 178–180].

At first, the biomass macromolecules are hydrolyzed into their respective monomers, whichare mostly soluble in water. These are then subject to dehydration and decarboxylationreactions, and intermediate products such as hydroxymethylfurfural (HMF) and furfuralare formed. The intermediate products then polymerize and form the biocoal (char) [81].Water, CO2, phenols and organic acids (acetic, formic, levulinic, propeonic, lactic acid)are byproducts formed by the decomposition of monomers and furfural-like intermediates[178]. Hexoses and pentoses, the monomers of cellulose and hemicellulose, respectively,have been found to undergo different reaction pathways [180]:

• cellulose → hexose (glucose, fructose) → HMF → char

• hemicellulose → pentose (xylose) → furfural → char

The reaction onset temperature of HTC is largely defined by the temperature at which hy-drolysis sets in, which is at approximately 180°C for hemicellulose and 200 °C for cellulose[81]. Dinjus reports that the formation of microspheres from cellulose starts at 210–220°C[158]. Lignin is only slightly affected in the usual temperature range of HTC. Lignin wasfound to remain as a porous matrix from which cellulose is dissolved away [158]. Biomasswith a high lignin content (such as wood) therefore retains its original shape during HTC.Experiments with lignin at 220°C and 4 h showed no difference in elemental compositionof the feedstock and the carbonization product [162]. However, Bobleter [27] suspects thatisolated lignin has undergone modifications during its extraction from the plant material,and that the lignin contained in plant materials is more accessible to degradation thanisolated lignin. Degradation products of lignin include monomeric phenols [27]. A com-prehensive review of the relevant reaction mechanisms participating in HTC is providedby Funke [181, page 31].

2.2.3.2 Products

In addition to the solid biocoal product, the reaction yields carbon dioxide and othergaseous products in small amounts and also water with soluble compounds dissolved.The solid phase has been the most thoroughly studied, in part because most of the earlyresearch effort was on natural coal formation, and most of the current research is aimedat using the biocoal as a soil conditioner. Indeed, complete mass and energy balances forall products are rarely reported. The distribution of carbon to the three product phases,however, is more commonly reported in literature. This data is summarized in Table 2.6.The mass yield of the solid product is heavily dependent on the operating conditions andis reported between 30–70%.

Attempts to formulate a stoichiometric equation for the net overall reaction can be foundin literature [148, 150]. Based on experimental data, Bergius [148] suggested Equations

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Table 2.6: Carbon yield of the three product phases from HTC [23, 163, 164, 182, 183].Values in brackets indicate experiments conducted in recirculated HTC process water.

T [°C] t [h] solid dissolved gasstraw, organic waste 81–84% 10–15% 4–6%green waste, organic waste 180 12 73–77% 16–21% 6%various organic waste materials 180–220 4–16 49–75% 16–46% 4–9%poplar wood 220 4 80% (86%) 16% (11%) 4% (6%)EFB 220 4 74% (77%) 18% (13%)

2.1 and 2.2 for the maximum possible carbonization of cellulose and lignin, respectively:

(C6H10O5)4 → C21H16O2 + 3 CO2 + 12 H2O (2.1)

C11H10O4 → C10H8O + CO2 + H2O (2.2)

These equations, however, do not take into account the dissolved products. CO2 andH2O are calculated by the differences in carbon, hydrogen and oxygen between biomassand biocoal, which leads to an overestimation of these products. Dinjus et al. formulatedstoichiometric equations with this approach and found that the CO2 calculated is typicallytwice that of the amount measured [158].

2.2.3.3 The solid product

In a similar manner to torrefaction, HTC removes oxygen from the feedstock by generatingCO2, H2O and dissolved compounds, thereby increasing the carbon content of the solid.The degree of carbonization depends on the reaction temperature and residence time.The higher the degree of carbonization, the lower the O/C ratio of the product, and thelower the mass and energy yields. The higher heating value of the biocoal increases withthe degree of carbonization — values of over 30 MJ/kg can be achieved with a reactiontemperature of 250°C [158]. Typically, HTC increases the higher heating value by 2–10MJ/kg (d.b.) compared with the feedstock biomass.

Figure 2.1 shows a Van Krevelen diagram of HTC biocoal from various feedstocks, torrefiedwood and straw, lignite and bituminous coal. H/C and O/C are the atomic ratios ofhydrogen to carbon and oxygen to carbon, respectively. The carbonization of wood andstraw mostly follows the path for cellulose according to Equation 2.1. For HTC of woodand straw at 220°C, the final product lies mostly in the range of lignite. The maximumpossible carbonization of cellulose and straw according to Equations 2.1 and 2.2 leadsto products in the range of bituminous coal. Experiments indicate that the maximumcarbonization of cellulose is reached after 64 h at 310°C [147]. Due to its high carboncontent, lignin requires relatively little carbonization to reach an O/C ratio similar tolignite. The analyzed empty fruit bunches from the palm oil production and old leaveshave a higher H/C ratio than the other feedstocks, and so has their biocoal. Compared totorrefied wood and straw, HTC biocoal is more severely carbonized, and therefore more

23


“coal-like” in its composition.

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wood 220°C straw, 200°C straw 220°C straw 240°C wood, various straw, variousgras, various torrefied willow torrefied strawtorrefied pine torrefied bagasse lignitebituminous coal poplar rec. 220°C, 240 min EFB, rec. 220°C, 2400 minleaves, rec. 220°C, 240 min cellulose, max. lignin, max.EFB, raw leaves, raw poplar, rawstraw, raw cellulose, raw lignin, raw

reaction direction

Figure 2.1: Van Krevelen diagram of experimental data for HTC biocoal and its feedstocks[23, 158, 163, 184, 185]. For comparison, data for the maximum carbonization of celluloseand lignin [148], torrefied materials [122, 128, 130, 141], lignite and bituminous coal[29, 186] are included. The term rec denotes experiments conducted in recirculatedprocess water. The lines connect the raw material and its respective biocoal for selectedfeedstocks.

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2.2.3.4 Gaseous byproducts

The gas phase (d.b.) obtained from the hydrothermal carbonization of cellulose at 225°Cwas found to contain 91% CO2, 6.4% CO, 0.8% H2 and 0.4% hydrocarbons [148]. The gasleaving the reactor is saturated with water vapour. The quantity of feedstock convertedinto gaseous products as well as the volume fractions of H2 and hydrocarbons increasewith reaction temperature. For temperatures up to 250°C, between 1% and 10% of thefeedstock carbon reacts to CO2 [23, 150, 152]. Mass yields of gaseous byproducts for HTCfrom poplar, old leaves and empty fruit bunches are shown in Table 2.7. The gaseousphase from the HTC of biodegradable waste was found to contain significant amounts ofH2S. 200 ppm of H2S were measured in the gas phase diluted with air [187], which maycorrespond to over 1300 ppm in the gaseous byproduct.

Table 2.7: Mass yield of gaseous byproducts from HTC of poplar, old leaves and oil palmempty fruit bunches (EFB), per kg of dry feedstock, at 220°C and 4 h [163]. The termrec denotes experiments conducted in recirculated process water.

poplar poplar rec leaves leaves rec EFB EFB recCO2 6.62% 10.1% 9.19% 10.22% 8.41% 9.31%CO 0.32% 0.5% 0.25% 0.274% 0.41% 0.75%CH4 0.0033% 0.004% 0.0039% 0.0031% 0.002% 0.003%H2 0.0025% 0.042% 0.0017% 0.0018% 0.0% 0.01%

2.2.3.5 Dissolved byproducts

The dissolved byproducts of HTC have yet to be completely characterized. The amountof feedstock carbon that ends up in the liquid phase can be determined via measurementsof the total organic carbon content (TOC) of the process water. The TOC depends on thewater/biomass ratio and whether fresh water or recirculated process water was used as thereaction medium. TOC values between 8 and 36 g/l have been reported [23, 162, 163, 183].

A number of substances present in the liquid phase have been identified by high perform-ance liquid chromatography (HPLC) analysis, but they account for only 36–71% of theTOC. The main products identified are acetic acid, glycolic acid and formic acid, withsmaller amounts of levulinic acid, glucose, hydroxymethylfurfural (HMF), furfural andphenol [23, 162, 163]. Table 2.8 shows the concentrations of dissolved products from theHTC of empty fruit bunches and poplar shavings. Gas chromatography–mass spectro-metry (GC-MS) also revealed aliphatic and aromatic hydrocarbons [183]. On total, themass yield of dissolved compounds per kg of dry biomass lies in the range of 5–30% [188].

Water is also formed as a byproduct of HTC. Since the amount of water produced issmall compared to the amount of water used as a reaction medium, its yield is hard tomeasure. The maximum carbonization of cellulose under Equation 2.1 suggests 33% ofthe dry feedstock is converted to H2O.

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Table 2.8: Concentrations of identified substances in the process water from the HTC ofoil palm empty fruit bunches (EFB) and poplar shavings at 220°C and 4 h [23, 162].The term rec denotes experiments conducted in recirculated process water.

feedstock EFB EFB rec poplardry biomass to water ratio 1:5 1:5 1:9acetic acid [g/l] 12.7 33.4 3.7glycolic acid [g/l] 1.9 4.1 3.1formic acid [g/l] 2.0 1.0 6.5levulinic acid [g/l] 0.2 1.3 0.6glucose [g/l] 0.3 0.9 0.9HMF [g/l] 0.2 1.3 1.4phenol [g/l] 0.8 0.0 0.4furfural [g/l] 0.3TOC of process water [g/l] 17.2 33.2 8.5

2.2.3.6 The fate of inorganic compounds

Some inorganic compounds are retained in the biocoal, while the remainder are dissolvedin the process water. In HTC experiments with dead leaves at 220°C, 77–88% of K, 70–78% of Mg, 22–46% of Na and 19–26% of N were dissolved in the liquid phase, while Caand P remained almost entirely in the solid product [163]. However, in HTC experimentswith EFB under the same operating conditions, more than 70% of Ca and P dissolved[23]. For low ash feedstocks, the ash content of the biocoal is usually lower than that ofthe raw biomass, while for high ash feedstocks, ash accumulates in the biocoal. Table 2.9shows the ash content of selected feedstocks and their respective biocoals.

Table 2.9: Ash content (d.b.) of feedstock and biocoal for HTC from wood [185], emptyfruit bunches [23] and digestate [159], and fraction of feedstock ash dissolved in theliquid phase.

feedstock biocoal dissolved ashwt% (d.b.) wt% (d.b.) [–]

wood 1.3% 0.3–0.7% 63–85%EFB 4.5% 3–3.6% 53–62%food waste 7.5% 11.2% 35%digestate 35.9% 55.8% 27%

Experiments with coconut fibre and eucalyptus leaves [189] indicate that approximately30–80% of sulphur and nitrogen remain in the biocoal, which leads to a slight accumulationof these substances. About 85% of the chlorine was removed from switch grass and cornstover during HTC at 230°C [190].

2.2.3.7 Heat of reaction

The HTC reaction is slightly exothermal. Funke and Ziegler report a heat of reactionfor cellulose of −1.07 MJ/kg (daf) and for wood of −0.76 MJ/kg (daf) measured with

26


differential calorimetry for HTC at 240°C [191]. Other publications report heats of re-actions 100–600% higher than those measured values, based on energy balances. Sincethese calculations only account for the feedstock and solid product, but neglect the dis-solved byproducts, they grossly overestimate the heat of reaction available for internalheat recovery within the process [191].

2.2.3.8 Influence of operating conditions

The degree of carbonization and the formation of gaseous and liquid byproducts depend onthe reaction conditions, the most important of which appear to be temperature, residencetime, and the biomass to water ratio.

The more oxygen that is removed from the feedstock, the higher the calorific value and thelower the energy yield. For practical applications, this means there is a trade-off betweenproduct quality (HHV) and quantity (energy yield). As shown in Figure 2.2, the energyyields from measured data are much lower than those calculated by the theoretical reactionequations for the carbonization of cellulose (Equation 2.1) and lignin (Equation 2.2) dueto the presence of dissolved components. This may be more pronounced in lab-scaleexperiments than in industrial HTC plants, because the experiments are often conductedwith a higher water to biomass ratio, and use distilled water rather than recirculatedprocess water. For a given HHV, the energetic yield is generally higher for wood thanfor straw. Since the calorific value of raw wood is higher than that of raw straw, strawneeds a stronger carbonization to achieve the same HHV. Data from torrefaction andbiomass steam processing (BSP) is also included in the diagram. Torrefaction resultsin a weaker carbonization, and therefore higher energy yields and lower heating values,than HTC. BSP, a form of torrefaction in superheated steam, achieves a similar degreeof carbonization to HTC with a higher energy yield. This may be due to the absence ofdissolved compounds. However, only two data points were available for BSP, and generalconclusions cannot be drawn regarding its performance relative to HTC and torrefaction.

Ruyter [153] suggested the following relation between degree of carbonization fHT C , tem-perature T in [K], and residence time t in [s]:

fHT C = 50t0.2e− 3500T (2.3)

This relation suggests that temperature and residence time are substitutable. To reachthe same degree of carbonization in half the time, the temperature has to be increased byapproximately 10°C. This formula can only be considered a rough approximation, not acorrect model of the reaction kinetics, because the reactions taking place at different tem-peratures are actually different. This can be seen in the composition of the byproducts.For example, the furfural content decreases with temperature, while guaiacol, a decom-position product of lignin, is only generated at temperatures above 230°C [192]. Whilelonger residence times generally lead to a stronger carbonization, which lowers the massyield, they also allow more polymerization [81], which increases the mass yield.

The relation between temperature, residence time and the molar O/C ratio of biocoalfrom beech wood, cellulose and lignin is shown in Figure 2.3. The diagram includes adata point for beech wood carbonized for 17 h at 228°C, which would be predicted with

27


Equation 2.3 using the measured data at 250°C and 4 h (both have the same value forf HTC ). The predicted value seems in good agreement with the measured values at 17 h.However, Figure 2.3 indicates that the carbonization of cellulose takes place in a narrowtemperature range and this behaviour is not captured by Equation 2.3.

The biomass to water ratio has an important effect on the solid and liquid yields. A highbiomass to water ratio leads to a high concentration of dissolved intermediates, which in-creases the chance of polymerization [81]. Solution equilibria may also limit the generation

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poplar, rec. 220°C, 240 min. poplar, 220°C, 240 min. EFB, rec. 220°C, 240 min.leaves, rec. 220°C, 240 min. wood 220°C wood, rawstraw 200°C straw 220°C straw 240°Cstraw, raw wood, various cellulose, variouslignin, various torrefaction, willow torrefaction, strawtorrefaction, pine BSP, straw cellulose, theoreticallignin, theoretical

Figure 2.2: Relation between energy yield and HHV (d.b.) for HTC [23, 149, 152, 163,184, 185], BSP [89] and torrefaction [122, 128, 130, 141]. The diagram includes thetheoretical conversion of cellulose and lignin (as lines) according to Equations 2.1 and2.2 for comparison.

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beech 17 h predicted

Figure 2.3: Molar O/C ratio of biocoal from beech wood, cellulose and lignin carbonizedfor 4 and 17 h at different temperatures. Data from [158].

of dissolved byproducts [81].

Since HTC is conducted in liquid water, the minimum required pressure is determinedby the saturation pressure at the reaction temperature. Beyond that, pressure was foundto have little influence [81]. However, Herrmann reports that a pressure 5–15 bar abovesaturation pressure does increase the carbon content of the product [167].

The HTC reaction does not require catalysts, but acidic conditions enhance the hydro-lysis of cellulose and the overall rate of reaction of HTC [81]. Since organic acids areformed during HTC, they can act as autocatalysts [81, 178]. However, based on heatrelease characteristics measured by differential calorimetry, Funke and Ziegler concludethat autocatalysis by organic acid products is unlikely [191]. Lynam et. al report that theaddition of calcium chloride or lithium chloride increases the energy yield by 10% and theHHV of the biocoal product by 2–3 MJ/kg. Since less than 10% of the salts remain in theproduct, recycling of the salt solution may be feasible [193].

HTC does not require comminution of the feedstock to fine particle sizes. Experimentswith saw dust, wood shavings and wood cubes with an edge length of 100 mm showed noinfluence of feedstock particle size on product quality [167].

2.2.3.9 Effects of process water recirculation

Many experiments show a rather low energy yield. This is in part due to the fact thatmost laboratory experiments are conducted in distilled water, and often with a high waterto biomass ratio. This leads to a greater proportion of organic compounds dissolvedin the water. In a commercial-scale HTC-plant, recirculated process water rather thanfresh water would be used for creating the biomass slurry. This improves heat recoveryand minimizes the amount of waste water to be treated, and reduces the fresh water

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requirement. Stemann and Ziegler [163] conducted experiments with recirculated processwater and de-ionized water for poplar wood which show that the recirculation of processwater has a significant effect on the energy and mass yields. While the TOC and theconcentration of organic acids in the liquid phase increase, the reactive components (HMF,furfural, glucose) do not. In experiments with empty fruit bunches, the TOC of the liquidphase was found to be 17.19 g/l when the experiment was conducted in de-ionized waterand 33.43 g/l when the HTC was carried out in liquid recycled from previous experiments.This procedure should approximate process water recirculation [23]. The HHV of thebiocoal increases by 0.6–1.0 MJ/kg [23, 163] when recirculated process water was used,perhaps due to an increased reaction severity caused by the lower pH value. Carboxylicacids produced during hydrolysis reactions are known to enhance the reactivity of someorganic compounds under hydrothermal conditions [79].

For HTC with poplar shavings, the total gas yield increases by 2 percentage points, andthe H2 yield is about 10 times higher relative to the experiments with de-ionized water,possibly due to dissociation of formic acid. The total carbon yield increases from 80% to86%, probably by enhanced polymerization of the intermediate products. While in theexperiments with de-ionized water, 16% of the feedstock carbon end up in the liquid phase,in the experiments with recirculated process water this is reduced to 11%. The reductionof dissolved compounds leads to an increase in energy yield by 6 percentage points [163].Another benefit from process water recirculation is that the dry matter content achievedby mechanical dewatering was 64%, compared to 55% when using de-ionized water [163].The only negative effect from process water recirculation is that the ash content of thebiocoal will increase.

2.2.3.10 Overall process design

The overall process includes the preheating and pressurizing of the biomass and water toreaction conditions, and the cooling and depressurization of the biocoal slurry. To producea useful upgraded biofuel for combustion, the biocoal also needs dewatering and drying,and possibly either milling to a defined particle size or pelletizing.

The published literature mostly focusses on the results of lab-scale experiments, and thechemistry of the HTC reaction itself. Several technology developers have released limitedinformation on their plant designs, energy balances and economic assessments. Most ofthis data is available from German language conference presentations or company websites[173, 174, 194–196]. Conversely, the scientific literature contains few examples of energybalances and economic assessments of the overall process. To the author’s knowledge,operating data from pilot plants has not been published in the scientific literature as ofApril 2012. Since lab-scale experiments are conducted in batch reactors, and industrial-scale processes often employ continuous designs, the respective yields and biocoal qualitymay deviate significantly. Blöhse, however, reports that data from a small pilot plant witha continuous flow reactor and a processing capacity of 20–30 t/h confirms the results ofprior batch experiments [197].

Stemann and Ziegler conducted a simulation study of a continuous HTC process witha semi-continuous biomass feeding system using lock hoppers. The energetic efficiency(HHV) is 74–78% [198]. Heinrich presents a simulation model of a HTC process with two

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batch reactors in time-delayed operation, and an additional vessel for the intermediatestorage of recovered steam. The energetic efficiency (HHV) is 81%, without consideringauxiliary units like waste water treatment and pelletizing [199]. With comparable assump-tions on losses and auxiliary energy consumption, a simulation model of a continuous HTCprocess with internal heat recovery indicates an energetic efficiency of 82–90% [20].7

The early literature on the hydrothermal upgrading of peat [154, 200, 201] provides valu-able insight into a process very similar to HTC of biomass. Although the properties ofHTC biocoal may vary from those of carbonized peat, it can be expected that many of theengineering challenges are similar. A technical report issued by Svensk Torvförädling in1960 (in Swedish) describes in detail the experience at operating a pilot plant for upgrad-ing peat. Problems encountered during the operation were mostly related to the difficultnature of the slurry, causing clogging, fouling, abrasion and corrosion of equipment andpipes [200]. Much effort was directed to the design of a suitable slurry pump and dewa-tering press [200]. However, the pilot plant was successfully operated for 10 years, clearlyestablishing that the process design was technically feasible.Funke and Ziegler [188] discuss the feasibility of the technical implementation of HTC,based on the literature about peat processing plants and equipment employed in otherbiomass processes where relevant to HTC. They identify the feeding system, heat recoverysystem and waste water as key areas requiring R&D effort.

2.2.3.11 Properties of HTC-biocoal as a fuel for combustion and gasification

An important property for combustion efficiency is the moisture content of the fuel. Al-though the biocoal is received from the HTC reaction as a slurry and could thereforebe considered “wetter” than the original biomass, there is a difference in the quality ofthe water accompanying the biocoal and the feedstock biomass. Unlike the raw biomass,the biocoal can be mechanically dewatered to a high degree, which is much less energyconsuming than thermal drying. Indeed, dry matter contents of 55–68% were achievedwith a laboratory press [23, 163]. The mechanical dewatering is facilitated by breakingthe cell structure and the destruction of hydrophilic oxygen containing functional groups.IR-spectra show that peaks for OH-groups decrease significantly through HTC [163]. Infact, some hydrothermal treatment technologies such as mechanical-thermal-dewatering[84, 202] and some of the early applications of peat upgrading targeted dewatering ratherthan carbonization [154]. A reduction in equilibrium moisture content indicates that thehydrophobicity is increased, which means that biocoal is more resistant to water damageand less prone to biological decay than the raw biomass [82].

Grinding biocoal to a particle size below 280 μm in a lab-scale mill required 135 kJ/kg,which is in the same range as torrefied wood. The particle size distribution after millingwas found to be similar to that of Rhenish lignite [203], indicating that biocoal can bemilled in conventional coal mills. The milled biocoal particles showed a more spherical,less fibrous shape compared to milled wood. This suggests that fluidization behaviour inpneumatic transport systems will be improved [203].While the HHV (d.b.) of biocoal is in the range of lignite, the volatile compounds fractionwith 64–69% [23, 203] is slightly higher, more similar to peat. Due to its similarity with

7[20] presents an earlier version of the plant design and simulation model analyzed in this work.

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lignite, combustion of biocoal in lignite burners should not be problematic. In combustionfacilities designed for bituminous coal, the higher volatile matter content may lead todevolatilization or incineration in drying and milling equipment. Operating temperaturesof the respective equipment may need to be reduced to avoid this danger [204].

Compared to torrefaction, HTC offers the additional benefit of washing out some prob-lematic inorganic compounds like chlorine and potassium. The decrease of the K/Ca ratioleads to an increase of the ash melting temperature. Experiments on the ash meltingbehaviour of biocoal from empty palm oil fruit bunches found the sintering temperatureincreased by 100–150°C, and the flow temperature by more than 300°C. The raw EFB hada low ash sintering temperature of below 1000°C. For dead leaves with a high sinteringtemperature of over 1300°C before treatment, the experimental results on the influence ofHTC remained inconclusive [205].

Although some problematic substances are washed out, their mass fraction in the biocoalmay be higher than in the raw biomass, if the relative mass loss in C, H and O exceeds therelative mass loss of the respective substance. For the HTC of coconut shells at 250°C,the nitrogen content increased from 0.90% to 0.98% and the sulphur content from 0.23%to 0.29% [189]. This accumulation may be more pronounced for feedstocks with higher Nand S contents.

Overall, the ash content of the biocoal is largely dependent on the feedstock (see Table 2.9).Biocoal from high ash feedstocks, such as digestate or sewage sludge, may require dedicatedcombustion units, as used in the cement industry for residue-derived-fuels.

Reza et al. analyzed the pellet properties of raw wood pellets, pellets of torrefied woodand pellets of HTC biocoal produced at a carbonization temperature of 200–260°C [142].They found that when fully immersed in water, HTC biocoal pellets remain intact for 15minutes to one week, depending on carbonization temperature. Pellets of raw pine woodand torrefied pine disintegrated after 0.5 and 5 minutes, respectively. Fungi growth onthe pellets was inhibited by HTC at temperatures of 230°C or higher. The abrasion indexwas also reduced by HTC. The higher the carbonization temperatures, the more resistantthe pellets became against abrasion. The mass density of HTC biocoal pellets was foundto be 2–30% higher than that of raw wood pellets, and the energy density 7–70% higher.These results indicate that HTC biocoal pellets will be of higher quality than torrefiedwood pellets or pellets from raw wood. However, it should be noted that Reza et al.investigated HTC biocoal pellets from pine wood, and the results may not be transferableto other feedstocks. Wood contains a high amount of lignin, which acts as a binder inpelletization. Biocoal produced from feedstocks with a lower lignin content may thereforedisplay a different pelletization behaviour and pellet properties.

Given the similarity to fossil coal, the most likely short-term application appears to beco-firing in coal-fired power stations. Pelletization will facilitate transport and storage andshould allow the biocoal to be processed together with fossil coal in existing coal mills andfuel handling equipment.

Besides large coal-fired power stations, industrial combustion equipment fired on pulver-ized lignite may provide an interesting application case for HTC biocoal. Pulverized ligniteis utilized as a cheaper alternative to natural gas in industrial drying facilities and CHPplants. It is milled and dried and is delivered with a moisture content of approximately

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10% and a particle size below 200 μm. The ash content of pulverized lignite is around4% [206, 207], biocoal with a higher ash content may therefore not be suitable for therespective combustion equipment.

HTC also offers promise as a pretreatment for entrained flow gasification. The destructionof the fibrous structure will probably avoid the problems encountered with pneumaticfeeding-systems when using raw biomass. Moreover, the energy demand for milling tothe required particle size is greatly reduced. Gasification experiments in a bench-scaleentrained-flow gasifier operated at atmospheric pressure showed biocoal and lignite tohave a similar gasification behaviour [203].

HTC biocoal is generally not suited as a substitute for wood pellets in domestic heatingapplications, since the quality requirements specified in DIN EN 14961-2 regarding theash content are unlikely to be met.

2.2.3.12 Biocoal as a soil amendment

Besides the replacement of fossil fuels, a second strategy for GHG mitigation using up-graded biomass is widely discussed: utilization as a soil amendment. Unlike raw biomass,carbonized biomass provides a stable form of carbon. The application of carbonized bio-mass to soil would therefore constitute a long term storage option for carbon previouslyabsorbed from the atmosphere by the original plants. Lehmann et al. estimate that 50%of the biomass carbon could be sequestered if the biomass was converted to biochar. Whenraw biomass is biologically decomposed, less than 20% of its carbon remains after 5–10years [156].

Besides serving as a carbon store, biocoal may achieve additional GHG mitigation by itspotentially positive effects on the soil, including higher biomass yields, decreased fertilizerdemand, avoided nutrient leaching and reduced GHG emissions from the soil. Stichnothesuggests using HTC biocoal as a substitute for peat in greenhouses. He concludes thatthis utilization scenario provides the highest environmental benefits relative to other usesof HTC biocoal [208].

In the context of soil application, carbonized biomass is usually referred to as biochar. Todistinguish between biochar obtained from pyrolysis and from hydrothermal carbonization,the terms pyrochar and hydrochar are sometimes used in the literature. Hydrochar as asoil conditioner is a relatively new field of research, whereas more data has been publishedon pyrochar.

The most important criteria when employing biochars as a form of carbon storage isits long term stability. While the residence time of pyrochar in the soil is estimated torange from hundreds to tens of thousands of years [209], that of hydrochar is considerablyshorter. Steinbeiss at al. calculated a mean residence time of 4–29 years for hydrocharfrom yeast and glucose [210]. These various estimates should not be taken as conclusive.

Soil structure and water and nutrient retention may potentially be improved by biocharapplication. However, this proposal still lacks a systematic investigation for different soiland biochar types. For example, it was found that pyrochar increases moisture retentionin some soil types but decreases it in others [209]. One fundamental difference betweenpyrochars and hydrochars is that the former are alkaline and the latter are acidic. However,

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Rillig at al. found that despite the acidic nature of hydrochar, the soil pH-value increased,due to microbial reduction reactions [211].

Pyrochar can potentially suppress nitrous oxide and CH4 emissions from soils and de-creases nutrient leaching [209]. Both effects have a beneficial impact on the GHG balance.Moreover, it may lead to the same crop yield with reduced fertilizer application [209],thereby reducing the embedded GHG emissions from the fertilizer production and spread-ing. Hydrochar was found to improve the plant availability of K and P, and reduce theplant availability of N. Regarding nitrogen, there may be a potentially beneficial bufferingeffect which prevents leaching [208].

Biochar has also been found to have detrimental effects on plants. Rillig et al. reporthydrochar inhibits plant growth, especially in higher concentrations, but may stimulatefungal symbionts [211]. Germination seems to be especially hindered by fresh hydrochar,perhaps due to certain volatile byproducts from the HTC absorbed by the biochar [192,208].

An interesting niche application is the application of biochar to contaminated soils, wherebythe adsorption of contaminants onto the char reduces their availability to plants. Pyrocharwas found to reduce the plant uptake of pesticide residues [212].

The desired properties of biochar for soil applications are often the diametric opposite ofthe properties required for combustion. While a hydrophobic behaviour is highly desir-able for a biofuel, a high moisture retention capacity will be targeted for use as a soilconditioner. The mineral compounds in biocoal should be minimized for combustion ap-plications, but nutrients should be maintained in biochar for soil enhancement. Thesedifferent desired properties may lead to quite different plant designs and operating para-meters, dependent on whether the production of biocoal is tailored to soil improvementor to the use as a biofuel.

All in all, the effects of biochar application on plants, soil ecology, and the GHG balanceseem complex. More research is clearly needed to understand and assess the possible risksand benefits. Soil applications of hydrochar will generally not be economically viable inGermany under present conditions [208]. While soil amendment may offer a long-termoption, it seems unlikely that it will be implemented on large scale in the near future.

2.2.3.13 Other biocoal applications

HTC is also discussed as a source for the production of specialized carbon materials. Onefield of research is the manufacture of carbon nanostructures, which could, for example,substitute carbon black or prompt the development of new materials for the constructionindustry [213]. If inorganic salts are added to the HTC process, hybrid carbon/metalmaterials can be produced [179]. Another possible HTC biocoal application is waterpurification, as an ion exchange resin or sorption coal [180, 214]. HTC coal has beenfound to be an effective absorbent for copper compounds from waste water [215].

HTC biocoal has also been tested as a fuel for an indirect carbon fuel cell in laboratorytrials, where it displayed more favourable properties than fossil coal [216].

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2.2.4 Anaerobic digestion

Anaerobic digestion is a biological degradation process by which organic matter is con-verted into a combustible gas, so-called biogas, consisting mostly of methane and CO2.The process is commercially applied as a treatment process for sewage sludge, municipalorganic waste and farm residues, and for biogas production from dedicated energy crops.The biogas can be combusted locally in gas or diesel engine CHP plants, or upgraded tonatural gas quality and fed into the natural gas grid. The utilization pathways dependheavily on the prevailing public policy measures. In countries where the policy supportfor renewable energy is focussed on electricity generation, such as Germany, the biogas ispredominantly used in CHP plants. In Sweden, by comparison, the gas is usually upgradedas a transport fuel or used for space heating [217].In 2011, there were 7100 agricultural biogas plants in operation in Germany, with a totalinstalled electrical capacity of 3500 MW [102, page 14]. Plant capacities are mostly in therange of 70–500 kWel [217]. The dominant substrates for biogas plants with electricityproduction in Germany are animal excrements (45%) and maize (35%) [217].One problem for biogas plants with integrated CHP units is the lack of heat demand atthe rural sites where biogas plants are typically located. About 30% of the thermal energyproduced by the CHP plant is required for heating the digester [28], but the remainderis often discharged. Decentral biogas plants in Germany utilize, on average, only 43% oftheir waste heat [217].The upgrading of biogas to natural gas quality, so-called biomethane, and feeding it intothe natural gas grid locally decouples heat and power production from the biomass con-version. In 2010, the total installed biomethane feed-in capacity in Germany amountedto 35000 m3

STP/h or 385 MWHHV [217]. Current energy policy aims to strongly expandthis conversion pathway. The German Act for Access to the Natural Gas Grid (GasNZV)contains the target of substituting 6 billion m3 (6%) of natural gas by biomethane by 2020and 10 billion m3 by 2030 [218].Upgrading the biogas to natural gas quality is expensive and only economically viable forlarge plants. The biomethane plants therefore usually have a much higher capacity thanbiogas plants with CHP units. The average feed-in capacity in 2010 amounted to 751m3

STP/h, or 8.3 MWHHV [219]. The larger feedstock demand of the biomethane plantsleads to an increased use of energy crops rather than animal excrements. The dominantfeedstock is maize silage (74%), while other energy crops include grass silage (7%), grain(3%) and maize crop mixes (6%) [217].The increasing use of maize as a feedstock is controversial, because of the adverse envir-onmental effects of maize monocultures, including groundwater contamination by nitrateleaching and herbicides, high energy input and low biodiversity [5]. If maize is used asthe feedstock, the crop production accounts for the majority of GHG emissions related tothe electricity production from biogas. Taking into account land use change effects, GHGemissions can be even higher than those from burning natural gas [45].Depending on the cropping system, the methane yield from grass silage can be similarto that of maize silage [35], but the biomass yield per hectare is lower [56] and the feed-stock supply cost are higher [35]. Farming practices which respect the requirements ofbiodiversity conservation lead to much lower methane yields [56].

35


Grass from extensively farmed grassland is considered as the feedstock for anaerobic di-gestion in this work. The analyzed conversion pathways cover biomethane production anda process for the co-generation of solid biofuel and biogas.

A second application of anaerobic digestion considered in this study is the treatment ofHTC waste water (section 4.5.13). Research on this topic is still in an early phase but thefirst laboratory-scale experiments show fast degradability and moderately high methaneyields [220].


Anaerobic digestion can be divided into four phases, carried out by different microorgan-isms: hydrolysis, acidogenesis, acetogenesis and methanation [221]. During hydrolysis, thebiomass polymers are broken up into smaller parts such as sugars, amino acids and fattyacids. During acidogenesis, these are further degraded to alcohols and lower chain organicacids, CO2 and H2. The organic acids are converted to acetic acid, CO2 and H2 duringacetogenesis. Finally, methanation takes place via two pathways: splitting of acetic acidinto CH4 and CO2, and the conversion of H2 and CO2 into CH4 and water [102]. Hy-drolysis is usually the rate limiting step for substrates difficult to degrade [221]. Sincethe metabolic products of one group of bacteria are the food for the next group, carefuldesign of the overall process is required in order to keep the system in a stable condition[102]. The microbial communities may contain several thousands of different varieties ofmicrobes, with interdependent activities. Due to this complexity, the anaerobic digestionprocess is not well understood in detail [222].

Buswell and Müller [223] present the complete degradation for a given biomass compositionas follows:

CaHbOc +(

a − b

4− c

2

)H2O →

(a

2+

b

8− c

4

)CH4 +

(a

2− b

8− +

c

4

)CO2 (2.4)

In practice, economic considerations limit the acceptable residence time in the digester,and the conversion remains incomplete [102]. Typical conversion rates for the organicmatter lie between 40–85% [35].

While sugar, starch and proteins are easily degradable, hemicellulose and cellulose re-quire longer residence times [102]. For lignocellulosic biomass, the lignin forming partof the cell walls shields the digestible substances from being hydrolyzed [224]. Ligno-cellulosic biomass is therefore ill suited for anaerobic digestion without a pretreatment.Pretreatment methods which dissolve the hemicellulose or make the cellulose accessiblefor hydrolytic enzymes by changing the lignin structure are being investigated [224]. Theyinclude thermal treatment by hot water or steam, chemical or thermochemical treatmentswith acids, alkalines or other additives [224, 225], and enzymatic treatment [226].

Increasingly detailed models of the reaction kinetics of anaerobic digestion have beendeveloped over the past 30 years, taking into account the different degradation processes,the microbial growth kinetics, and inhibition mechanisms. An overview is given in [227].

36


2.2.4.2 Operating conditions and digester design

An important parameter for the reactor design is the organic loading rate (OLR), definedas the daily amount of organic dry matter moDM fed into the reactor per m3 of activedigester volume. The OLR is typically below 4 kgoDM/m3/d. If the OLR is too high, thesystem becomes overloaded, because more substrate is added than the microbes can digest[102]. The residence time in the digester is referred to as hydraulic retention time (HRT)and is directly related to the OLR

HRT =Vfermenter

Vfeed

=ρ · xoDM

OLR(2.5)

where Vfermenter is the active digester volume, Vfermenter is the daily feed inflow in [m3/d],ρ is the mass density of the digester content and xoDM the organic matter content of thefeed. Since most biogas is produced within the first 30 days, there is a trade-off betweengas yield per digester volume and gas yield per mass of substrate organic matter. TypicalHRT lie between 30 and 120 days, depending on the degradability of the substrate [102].

Higher operating temperatures require lower residence times for the same conversion rate[102]. However, in the thermophilic temperature range between 50–55°C, the process isless robust than in the mesophilic temperature range between 38–42°C, due to a lowerdiversity in microbes. A temperature drop of little more than 1°C for a few hours can leadto a disturbance of the microbial community with a weeks long recovery phase [102].

Other operating parameters such as pH-value, nutrient supply, and substances which aretoxic for the microbes or inhibit the biogas production need to be carefully monitoredand controlled. Substances which may enter the digester with the substrate and inhibitthe biogas production include heavy metals, oxygen, antibiotics, too much nitrogen, andmouldy feedstock [102].

A number of different reactor designs are employed for anaerobic digesters. One importantcriteria is whether the process is conducted with wet pumpable substrate with a dry mattercontent of 5–15% or with so-called dry stackable substrate with a dry matter content ofmore than 20%. Furthermore, the digester can be designed as a continuously stirred tankreactor or as a plug flow reactor, digestion can be performed in one or several stages, andfeeding can be in batch or continuous mode. The most commonly used configuration foragricultural biogas plants is a continuously stirred tank reactor operated at mesophilictemperatures processing wet pumpable substrate with continuous feeding [28, 102]. Thisprocess design has been developed for liquid manure. With the use of energy crops as asubstrate on the rise, there is increasing interest in the development of alternative reactordesigns optimized for feedstocks with a higher dry matter content [228].

Anaerobic digestion is exothermal, but, due to heat losses, the digester needs to be heated.

2.2.4.3 Products

The products of anaerobic digestion are the biogas and the digester residue. This residue,referred to as digestate, contains the unconverted organic material and the inorganic sub-stances.

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The biogas consists mostly of CH4 and CO2, as well as small amounts of H2, H2S and NH3[229]. Methane yields, methane fraction of the biogas, conversion rate of the dry organicmatter and energy yield for different feedstocks are summarized in Table 2.10.

The methane yield for grass strongly depends on the time of harvest. Grass becomesmore lignocellulosic with age, resulting in a worse digestibility and decreasing methaneyield. While from young grass, 70% of its energy can be recovered in the biogas, for oldgrass, this reduces to 15% [221]. This typically results in low yields where biodiversityconservation necessitates a late cut [56].

Table 2.10: Biogas yields and conversion efficiencies for anaerobic digestion using differentfeedstocks, based on [35, 56, 230].

substrate OLR CH4

yieldCH4

contentconversion

rateenergyyield

(HHV)[kg/m3/d] [l/kgoDM] [vol%] [% oDM] [–]

maize silage 4 338 52% 85% 67%forage beet and leaf 456 90%organic waste 4 369 60% 76% 73%rye grass 410 81%grass silage 4 318 53% 79% 63%fresh grass, untreated 3 324 54% 78% 64%grass, park and gardening waste 3 150 50% 41% 30%grass, intensive cropping 350 69%grass, biodiversity grassland 230 46%

The digestate is mostly applied to agricultural land as a fertilizer [231]. In order to avoidmethane emissions during after-fermentation, it is mandatory to store the digestate in agas-tight residue storage tank for some time before it can be spread. The residue storagetank needs to be large enough to hold the digestate during the time when the soil is frozenand spreading of the digestate is prohibited [221]. The amount of digestate which can beapplied to a given area is limited by ground water protection legislation and maximumallowable annual nitrogen loads [221, 231]. For large biogas plants, the land area requiredto distribute the digestate results in long transportation distances. Treatment of thedigestate is therefore attracting increasing interest [231].

Solid-liquid separation by screw press, filter press or decanter is commonly applied [221,231]. The press water is mostly used as a fertilizer [28], while the thickened phase witha solid matter content of 20–40% is mostly composted [221]. Further treatment to purifythe liquid phase or produce standardized fertilizers are currently not common due to highcosts but may become increasingly important [231].

In addition to causing a disposal problem, the unconverted organic matter also limits theenergy yield of the anaerobic digestion process.

38


2.2.4.4 Biomethane production

If biogas is to be fed into the natural gas grid, it has to be equivalent in quality to naturalgas at the point of feed-in. The requirements regarding composition, higher heating valueand dew point are therefore considerably higher than for the combustion in an on-siteCHP plant. In particular, CO2 removal is necessary to reduce the CO2 content from over40% to the maximum allowed value of 6% [232]. The most commonly used technologiesfor CO2 removal are pressurized water scrubbing and pressure-swing adsorption (PSA).Other options include absorption with solvents such as amines, glycol or selexol [28].

Depending on the CO2 removal process, upstream desulphurization and/or downstreamdrying of the gas is required. Pressurized water scrubbing has the advantage that itremoves both CO2 and H2S and no prior desulphurization is required. Drying can beaccomplished by cooling and removing the condensate, or by adsorption using zeolithes,silica gel or aluminium oxide [232].

2.2.4.5 Integrated generation of solid biofuel and biogas

Richter et al. suggest a process whereby herbaceous biomass is separated into a pressfluid and press cake using a screw press. The easily digestible press fluid is then convertedto biogas by anaerobic digestion, while the fibrous press cake is dried for utilization as asolid biofuel [233, 234]. The biogas is combusted in a reciprocating engine for electricitygeneration, while the engine waste heat is utilized to dry the press cake. Richter et al. referto the process as Integrated generation of solid Fuel and Biogas from Biomass (IFBB).

In batch trials conducted under mesophilic conditions with grass silage from conservationgrassland, 90% of the organic dry matter of the press juice was converted after 13 days[233]. Methane yields were 410–520 l/kgoDM [234]. On the other hand, anaerobic digestionof the whole crop silage returned very low methane yields of 173-285 l/kgoDM [234], and aconversion rate of less than 60% after 27 days [233].

Hydrothermal conditioning of the substrate for about 15 minutes at 60–80°C before mech-anical separation enhances the mass flows into the press fluid [235]. Over-proportionalshares of P, K, Mg, S, Cl and, to a lesser extend, nitrogen, end up in the press fluid.This is advantageous for two reasons: firstly, less K and Cl in the press cake improve itsquality as a combustion fuel. Secondly, more P, K and N in the press fluid means thatmore nutrients can be returned to the field with the digestate [234, 235].

The methane yield per mass of grass silage is lower for the IFBB process than for the whole-crop digestion, but the total energy yield is higher because of the additional productionof the solid biofuel [53, 233, 234]. Additionally, the problems with undesirable byproductsare resolved: there is only a small amount of digestate, and waste heat from the associatedCHP plant can be utilized for drying the press cake [234].

The digestion of the press fluid, which has a low dry matter content of below 2%, can bebased on digester techniques developed for waste water treatment plants [233].

The overall IFBB process comprises the unit operations hydrothermal conditioning, mech-anical separation, anaerobic digestion of the press fluid, combustion of the biogas in a CHPplant, and drying of the press cake [233].

39


2.3 Bioenergy with carbon capture (BECCS)

When sustainably grown biomass is combusted in a facility with carbon capture, the CO2previously taken up by the plants from the atmosphere can be deposited in a storage site,resulting in net negative CO2 emissions.

Generally speaking, the same carbon capture technologies developed for fossil fuels canbe employed for biomass. Since carbon capture technology is likely to be developed forcoal in the first instance, biomass upgrading technologies which make the biomass more“coal-like” may facilitate the combination of bioenergy with CCS. The two key candidatesfor use with BECCS are, therefore, torrefaction and HTC.

The following section provides an overview of CCS in general. Gasification with pre-combustion carbon capture, the technology which is analyzed in this work, is described insome detail in section 2.3.2. The potential role for biomass upgrading in the context ofgasification with CCS is discussed in section 2.3.3.

2.3.1 CCS technologies

CCS consists of the three steps: (1) capture, i.e. the retrieval of the CO2 at the powerplant or other industrial facility, (2) transport to the storage site and (3) long-term storage.

Carbon capture technologies for power stations can be classified based on the point atwhich the CO2 separation occurs within the fuel conversion process. In post-combustioncapture, the CO2 is separated from the flue gas after the combustion. The most commonlyproposed separation technology is absorption with a chemical solvent, such as amines.The combustion process itself does not need to be altered for post-combustion capture. Itcan therefore be retrofitted to existing power stations and is suitable for both solid andgaseous fuels. The main disadvantage of post-combustion capture is the fact that the CO2in the combustion gases is diluted with large amounts of nitrogen from the combustionair, leading to a high energy penalty for separation.

Pre-combustion carbon capture is the decarbonization of the fuel prior to the combustion.The most discussed pre-combustion capture technology is gasification of a solid fuel fol-lowed by a water gas shift reaction and subsequent separation of the CO2 by absorptionwith physical solvents. Due to the high partial pressure of CO2 in the product gas ofthe shift reactor, the energy demand for separation is significantly lower than for post-combustion capture. Pre-combustion decarbonization of natural gas can be achieved bymethane steam reforming followed by the water gas shift reaction and CO2 absorption.

Research on CO2 separation for post- and pre-combustion capture focusses on the devel-opment of better solvents as well as work on alternative separation methods includingsolid sorbents such as CaO, pressure swing adsorption with molecular sieves or activatedcarbon, and membrane processes [236].

The third type of capture is referred to as oxyfuel combustion and denotes combustion inoxygen instead of air. The combustion products consist mostly of CO2 and water, thusthe CO2 can be readily separated by cooling and condensing the water. Air separation toprovide the oxygen is energy intensive. Several concepts which avoid the necessity for airseparation have therefore been suggested in literature. One example is chemical looping

40


combustion, where the oxygen is delivered to the fuel by a solid oxygen carrier, usually ametal oxide. The reduced metal is then re-oxidized with air in a separate reactor.

Some industrial processes require the removal of CO2 from a material stream. While todaythis CO2 is mostly vented to the atmosphere, CCS could be employed. Such processesinclude the sweetening of natural gas, where CO2 is removed by chemical or physicalabsorption to fulfil quality requirements. The production of biomethane via anaerobicdigestion of biomass also requires the removal of CO2 for feed-in into the natural gas grid.However, since most anaerobic digestion plants are small-scale in rural areas, they maynot have easy access to any future CO2 transport infrastructure.

The transport of large quantities of CO2 is likely to be conducted by pipeline. This isa mature technology which is employed for enhanced oil recovery. Since the CO2 is besttransported in supercritical state, prior compression to pressures above 80 bar is required.In order to avoid corrosion of the pipelines, hydrogen sulphide and moisture must beremoved. For small quantities of CO2 and transport distances over say 1000 km, shippingof liquefied CO2 may be less costly than pipeline transport [236].

Potential sites for geological storage comprise saline formations, gas and oil fields andunminable coal beds. Experience with large-scale geological CO2 storage is so far limitedto three projects where CO2 from natural gas processing is stored in saline formations,namely Sleipner and Snøhvit in Norway and In Salah in Algeria. Sleipner, the oldest ofthe three, came into operation in 1996 [237]. Some projects involving the injection of CO2for enhanced oil recovery have been in operation since the 1970s and 1980s [237].

There are currently no large-scale power stations with carbon capture in operation, butseveral pilot plants have been installed in the last few years. A pilot plant with a 30MWth oxyfuel boiler has been in operation at the power plant Schwarze Pumpe since2008, comprising burners for pre-dried lignite, exhaust gas cleaning, CO2 separation andliquefaction, and an air separation unit [238]. Combustion tests with bituminous coal havebeen successfully conducted with a 40 MWth oxyfuel boiler in Renfrew, UK [239]. At theWillem-Alexander IGCC plant in Buggenum, Netherlands, a slip stream of the syngasproduced from coal and biomass is provided to a pilot plant for pre-combustion carboncapture. The project started operation in 2011 [240]. Projects with post-combustioncapture include a pilot-plant at the lignite-fired power station in Niederaußem, Germany[241]. An overview of international operating and planned CCS projects is provided in[242].

Numerous simulation studies have been presented in literature which assess the impact ofcarbon capture on the efficiency and costs of coal and natural gas fired power stations. Thedecrease in efficiency caused by the energy demand for the separation and compressionof the CO2 lies in the range of 7–11 percentage points in most studies, resulting in anefficiency of 28–36% for a coal-fired power station with CCS. The additional investmentamounts to 500–900 €/kWel for a natural gas-fired combined cycle plant and 900–1700€/kWel for a pulverized coal-fired power station [243–246]. Regarding efficiency and cost,there is no clear advantage for either post-combustion capture, pre-combustion captureor oxyfuel combustion for coal-fired power stations. Finkenrath concludes, based on datacompiled from several studies, that the differences in the cost of electricity between thethree capture routes lie within the range of uncertainty of feasibility study cost estimates.None of the three capture routes can therefore be considered economically superior [243].

41


Capture rates between 85% and 100% of the feedstock carbon are reported in literature,being higher for oxyfuel combustion than for post- and pre-combustion capture [243].Gasification with pre-combustion capture offers the advantage that the syngas can beused for electricity production as well as the production of second generation transportfuels via Fischer-Tropsch synthesis (FTS). On the downside, it is the most complex of thethree technologies.

Relatively few studies have analyzed the cost and efficiency penalties of carbon capturefor power plants fuelled by biomass, although the general concept of BECCS has recentlygained attention in global energy system models. Domenichini et al. report an efficiencypenalty of 14.5 percentage points for a 170 MWel plant with a fluidized bed boiler firedon wood and post-combustion carbon capture. This leaves the BECCS plant with anefficiency of only 23.6% (HHV)8 [247]. Rhodes and Keith report efficiencies of 25–28%(HHV) for biomass-fired IGCC plants with pre-combustion capture and a net electricitygeneration of 110–123 MWel [248].

2.3.2 Gasification and pre-combustion carbon capture

A simplified flow diagram of gasification with pre-combustion carbon capture is shown inFigure 2.4. The raw gas consisting of CO, H2, CO2, H2O, CH4 and other hydrocarbons aswell as various contaminants is cooled down to the temperatures required by the respectivecleaning steps. Particles, alkali metals, sulphur and chlorine components and tar areremoved in a series of gas cleaning steps which are discussed in section 2.3.2.2. CO isconverted to CO2 with the water gas shift reaction and separated from the syngas, mostlywith physical absorption. The clean syngas consists mainly of hydrogen. In the followingsection, different types of gasifiers for coal and biomass and their compatibility with thisgeneral process design are discussed.

gasificationgas cooling& cleaning

COseparation

2water gas shiftCO+H O CO +H2 2 2�

ash

steamgasification agent (air, O , H O)2 2

particles,alkali metals,

sulfur,tar,...

CO2

feedraw gas clean gas

CO, H ,C H ,...

2

x y

mostly H2

Figure 2.4: Simplified flow diagram of gasification with pre-combustion carbon capture.

2.3.2.1 Gasification

Numerous gasification technologies have been developed. Roughly, they can be classifiedbased on the following characteristics:

8Estimated from 25.8% (LHV).

42


• gasifier type: entrained flow, fluidized-bed or fixed-bed• feed delivery to the gasifier: slurrified or dry• operating pressure: atmospheric or pressurized• gasification agent: oxygen, steam or air• ash removal: dry or as melted slag

For bituminous coal gasification, with or without CCS, the preferred technology is pres-surized oxygen-blown entrained flow gasification. Due to the high operating temperatures,usually above 1400°C [249], entrained flow gasification produces a tar-free syngas with avery low methane content and has a carbon conversion of over 99% [249, page 121]. Oper-ating pressures are in the range of 20–70 bar [249, page 120]. The syngas consists mostlyof H2 and CO, which makes it well suited for pre-combustion carbon capture and for theproduction of transport fuels with FTS.The short residence time9 results in compact gasifier equipment, but requires the feedstockto be provided at small particle size (0.1 mm for coal). Some gasifiers, like the ShellCoal Gasification Process (SCGP) and the Prenflo process, use dry feedstock pressurizedpneumatically by lock hoppers, while others, such as the GE Energy process (formerlyTexaco), are supplied with a pumpable coal/water slurry. While the later is technicallyeasier, it decreases the efficiency because large quantities of water are introduced into thegasifier.Since the gasifiers are operated above the ash melting temperature, the ash is generallymelted and removed as slag. A disadvantage of the high operating temperature is the highoxygen demand, especially for slurrified feed or feedstocks with a high ash content [249,page 120]. This leads to a high auxiliary energy consumption for the air separation.Several large-scale coal-fed entrained flow gasifiers with a capacity of several hundredMWth are in operation. The dominant use of the syngas is for the production of chemicalssuch as FTS-fuels, ammonia, methanol and hydrogen. In the last years, new plants havemostly been built in China [250].For raw biomass, however, pressurized entrained flow gasification is infeasible becauseof problems related to the feeding system. Laboratory-scale tests with pulverized woodindicate that a pneumatic feeding system is not suitable for wood powder, due to the highcohesion between the fibrous particles [251].Preparation of a biomass-water slurry is equally infeasible, due to the hygroscopic natureof the biomass. Based on viscosity measurements, the maximum solids content for apumpable slurry with pine wood particles of 0-150 μm was found to be below 13%, whilethat of a sub-bituminous coal was 65% [252]. Liquid organic waste such as glycerol frombiodiesel production, waste alcohols or liquid waste compounds from biorefineries havebeen suggested as fluids for creating a pumpable biomass slurry [253]. However, thepotential capacity of biomass gasification would then be limited by the availability ofthese substances.For medium-to-large scale biomass gasification, air-blown atmospheric fluidized bed gasi-fiers are most common. Both bubbling and circulating fluidized beds are employed [250].Typical plant input capacities range up to 50 MWth [28, page 613].

9Residence times are in the range of 0.5–4 sec for the SCGP and the Prenflo process [249, page 131].

43


Fluidized bed gasifiers accept a particle size of up to 70 mm [28, page 611] and wood isthe prevalent feedstock.

Operating temperatures lie mostly in the range of 700–900°C [28, page 609]. The loweroperating temperature compared to entrained flow gasification results in a higher CH4content in the syngas, as well as other hydrocarbons and tar. While tar is a much discussedproblem in biomass gasification, methane and other light hydrocarbons do not presenta problem in IGCC applications without carbon capture. They have a high calorificvalue and are readily burned in the gas turbine. However, in a process scheme with pre-combustion carbon capture according to Figure 2.4, the methane would pass the watergas shift reactor unconverted, therefore a significant share of the feedstock carbon wouldevade the capture.

The higher operating temperatures required to reduce tar and methane content are gen-erally not possible, because sintering of the ash would destroy the fluidized bed. Wanget al. assess ash related problems including sintering and corrosion to be one of the keyissues of biomass gasification [254]. Sulphates and alkali silicates formed during the gas-ification with melting points below 700°C can confine the operating temperature to evenlower values [254]. Herbaceous biomass with its high ash and alkali content is far moreprone to ash-related problems than wood [255]. This may partly explain why gasificationof herbaceous biomass is in a much less mature state of development.

Air-blown gasification leads to dilution of the syngas with nitrogen, which makes it poorlysuited for carbon capture. To improve the quality of the syngas, gasification with steamis widely discussed for biomass. Since steam gasification is endothermic, thermal energyhas to be provided to the gasifier. To this end, char can be extracted from the gasifierand burned in a second fluidized-bed reactor. Thermal energy is transferred to the gasifierby the exchange of bed material between the two reactors [28, page 614]. This concepthowever is not well suited to pre-combustion carbon capture, because part of the feedstockcarbon is burned rather than gasified and cannot then be captured.

Pressurized gasification with oxygen or an oxygen-steam mixture seems to be the mostsuitable concept for a biomass-fired IGCC with pre-combustion carbon capture as well asfor chemicals production. Ciferno and Marano in 2002 [256, page 18] analyzed biomassgasification technologies in terms of their suitability for FTS, methanol and hydrogen pro-duction. Of the 15 gasifiers covered by their survey, the RENUGAS pilot plant10 was theonly one operated at a pressure above 20 bar with oxygen-steam mixture. Kirkels andVerbong (2011) conclude that “the stage of development of biomass gasification can bestbe characterized as one of limited niche development” [250]. They identify scaling up andfurther development of tar reduction and gas cleaning as crucial for a large-scale mar-ket entry. Three oxygen-blown pressurized gasifiers11 are amongst their most promisingbiomass gasification concepts.

10The RENUGAS plant was developed by the Gas Technology Institute (GTI), with a capacity of 136–455kg/h (<2.5 MWth) and is no longer in operation [256, page 37].

11RENUGAS, Siemens Fuel Gasification Technologies (Future Energy), and BIOSYN technology by En-erkem Technologies Inc.

44


2.3.2.2 Syngas conditioning

For IGCC applications without CCS, the main objective of the syngas conditioning is theremoval of contaminants which could damage downstream equipment or which must notbe released to the environment. These comprise particles, alkali metals, sulphur (H2S),chlorine (HCl), nitrogen compounds (NH3, HCN), and tar. For the synthesis of FTS-fuels or methanol, there is the additional requirement of a low hydrocarbon content anda specified H2/CO ratio [256, page 13].Particles of ash, char, and bed material cause erosion and fouling of downstream equip-ment. They can be removed with the help of dry cleaning technologies including cyclonesand various types of filters, or wet cleaning using scrubbers or electrostatic precipitators[257].Alkali metals, mostly sodium (Na) and potassium (K) form compounds which evaporateat temperatures above 800°C during gasification. They contribute to corrosion and canbuild deposits on plant equipment when the raw gas is cooled down to below 600°C [28,page 626]. If the alkali content exceeds the acceptable level for the gas turbine, it canbe reduced by adsorption with bauxite or other sorbents. In fluidized bed gasification,the sorbents can be mixed into the gasifier bed, otherwise a downstream sorbent bed isrequired [83]. Below 500°C, alkali compounds are solid and can be removed by particulateremoval methods [249, page 308]. For feedstocks with K contents below 0.17 kg/ GJfuel,the fouling potential is low [83]. The K content of short rotation poplar and willow is0.13–0.18 kg/GJfuel

12, alkali removal therefore may or may not be required. Catalystpoisoning can be another consequence of alkali compounds [256, page 10]. This may setstricter limits on the alkali content depending on the downstream plant components.Chlorine can cause corrosion and catalyst poisoning [256, page 10]. Since the chlorinecontent of wood is very low, chlorine removal is often not required when untreated woodis used as a feedstock [28, page 640]. Coal and herbaceous biomass have a significantlyhigher chlorine content and require chlorine removal, mostly accomplished by scrubbing[28, page 640]. High temperature chlorine removal by adsorption to sodium carbonate isanother option [258, page 28].Removal of nitrogen compounds (NH3, HCN) is usually required due to statutory limitson NOx emissions from the gas turbine exhaust gas [28, page 627]. These are usuallyremoved by scrubbing, but catalytic conversion at around 900°C with dolomite or nickelis under development for biomass gasification [28, page 639].Issues related to sulphur, mainly present as H2S, comprise corrosion, catalyst poisoning,and emissions [256, page 10]. The sulphur content of untreated wood is generally verylow [28, page 627]. For gases with a low H2S concentration, adsorption in a ZnO-bedat 350–450°C is a commercially proven technology [28, page 639]. In coal gasification,where the H2S concentration is much higher, absorption with solvent regeneration is thestandard desulphurization technology. Solvents include amines, rectisol (cold methanol)and selexol (dimethyl ethers of polyethylene glycol, DMPEG) [249]. One disadvantage ofthese processes is their low operating temperature of below 40°C.Condensable hydrocarbons, referred to as tars, remain one of the main challenges of bio-mass gasification [83, 257]. When the syngas is cooled down or compressed, they form12poplar: 0.35% K, HHV=19.8 MJ/kg; willow: 0.26% K, HHV=19.7 MJ/kg [28, pages 343, 360]

45


deposits on plant equipment [28, page 625]. Primary tar reduction methods attempt toavoid tar formation or to convert tar within the gasifier. They include the selection ofoperating conditions (with temperature being strongly influential), gasifier design and theuse of bed additives [259]. Secondary methods refer to gas treatment downstream of thegasifier and either aim to convert the tars to unproblematic, shorter-chained gas compon-ents, or to separate the tars from the gas stream. The former comprise thermal crackingat temperatures above 1000°C [28, page 638] and catalytic cracking with dolomite, olivine,alkali metals, or nickel-based catalysts, which takes place at 800–950°C [28, 257, page 638].The later include scrubbing in water or organic solvents, and electrostatic precipitation.The use of water for tar removal can result in large amounts of contaminated water, whichis difficult to treat [28, page 636]. Since tars have a high calorific value, their removal mayimply a significant energy loss [260]. Tar removal methods for biomass gasification andthe influence of the gasifier operating conditions on tar formation are extensively coveredin the literature. Overviews are presented, for example, in [83, 254, 257–259].

It is noteworthy that, for IGCC applications, tar removal is not necessarily required. Ifcooling of the syngas below the condensation temperature of the tars is avoided, theyare simply burned in the gas turbine combustion chamber [28, page 631]. The biomassgasification IGCC pilot-plant in Värnamo, Sweden, successfully operated a candle filter at350-400°C which let the gaseous tars pass through while the condensed tars were removedby the filter [261].

For IGCC applications with carbon capture, at least two steps are added to the syngasconditioning: the water gas shift reaction and the separation of CO2.

CO is catalytically converted to CO2 with the water gas shift reaction at temperaturesbetween 200–500°C. Various catalysts can be employed, including iron oxide-based andcopper-zinc-aluminium compounds. While some of these are highly sensitive to sulphur, acobalt-molybdenum catalyst requires sulphur to maintain its active state [249, pages 348-351]. This so-called “raw shift” or “sour shift” is especially well suited for coal gasificationwith carbon capture, since desulphurization, which operates at lower temperatures, cantake place downstream of the shift reactor together with the CO2 separation.

The CO2 separation is mostly realized by absorption. Potential solvents are largely thesame as those for H2S removal.

When the syngas contains high levels of CH4, an additional methane steam reformingstep upstream of the shift reactor may be required to achieve a high carbon capture rate.Commercial nickel-based steam reforming catalysts have been shown to reduce the CH4content of syngas from steam-oxygen-blown biomass gasification by 87–99% in laboratoryscale trials. Tar was also eliminated in the catalytic bed which was operated at around800°C [262].

The design of a coherent syngas treatment concept for gasification is a complex issue. Theremoval of all relevant contaminants generally requires several cleaning steps, which mustbe matched to each other. For example, the use of catalysts in certain cleaning steps maypose additional requirements on upstream cleaning processes in order to avoid catalystpoisoning. Unnecessary cooling (and reheating) of the syngas should be avoided in theinterests of energetic efficiency and cost. This needs to be considered for the combinationof several cleaning processes, which all are limited to certain temperature ranges.

46


Generally, syngas treatment concepts can be subdivided into hot gas cleaning and cold gascleaning. Cold gas cleaning usually involves wet scrubbing and H2S removal by absorptionat temperatures below 100°C. Hot gas desulphurization, operating between 200–500°C, haslong been the subject of extensive research efforts, since it would eliminate the need forgas cooling to such low temperatures. However, for coal gasification, it remains an R&Dexercise [249, page 346].

Since most contaminants except for sulphur can be removed by dry cleaning technologiesat higher temperatures, hot gas cleaning seems to be more feasible for biomass gasification,due to the low sulphur content of most biomass.

2.3.2.3 Overall IGCC design

Apart from the gasifier and syngas conditioning section, an IGCC comprises an air separa-tion unit (ASU) if oxygen is used as a gasification agent, a sulphur recovery unit (typicallya Claus/SCOT plant), and the combined cycle power plant section.

Since the 1990s, there are several coal-fired demonstration plants in operation, mostlywith a capacity of 250–300 MWel [249, page 7].

With pre-combustion carbon capture, the hydrogen content of the syngas is considerablyhigher (up to 90%), which may cause problems in the gas turbine system with prematureignition, flame stability, combustion noise and emissions [263]. Experience with high-hydrogen content gas turbine fuels exists mostly from burning refinery gas [264]. OngoingR&D activities of gas turbine manufacturers aim at adjusting a range of gas turbines tosyngas applications with and without carbon capture [264, 265].

In some IGCC schemes, the gas turbine system is integrated with the air separation unit.The ASU is, in this case, operated at a higher pressure and air is provided to it fromthe gas turbine compressor. A high degree of integration may increase the efficiency butcauses difficulties regarding start-up and overall control. Newer plant designs thereforeseek a lower degree of integration, whereat about 30% of the ASU air is provided by thegas turbine compressor and the remainder by a separate compressor [249, page 295].

The first and (to the author’s knowledge) only biomass-fired IGCC was in operation from1993–1999 in Värnamo, Sweden. The 18 MWfuel plant comprised an air-blown, pressur-ized fluidized-bed gasifier, hot gas clean-up with a candle filter, and a combined cycledelivering 6 MW of electricity and 9 MW of district heat [261]. As part of the CHRISGASproject spanning 2006–2010, the demonstration plant was to be retrofitted to make highquality syngas for the production of synthetic transport fuels. The new design comprisedgasification with an oxygen-steam mixture, hot gas cleaning, methane steam reformingand a water gas shift reactor. However, this upgrade stalled through lack of funding. Gas-ification and syngas conditioning — including a hot gas ceramic filter operated at 600 toover 800°C, methane steam reforming and water gas shift — was instead analyzed in twopilot plant configurations of 10 kWfuel and 100 kWfuel [266]. The findings of this projectprovide useful insight for the design of biomass-fired IGCC with CCS, since most of thegas conditioning steps are identical.

47


2.3.3 Technology options for BECCS

As discussed in the previous section, entrained flow gasification is the most suitable incombination with CCS, but is currently not feasible for biomass due to feeding systemproblems. This leaves three options for biomass gasification with CCS:

1. adopt fluidized bed gasification and condition the syngas to enable a high carboncapture efficiency,

2. develop a dedicated biomass feeding system for entrained flow gasifiers,

3. pretreat the biomass to make it suitable for fine milling and use in the conventionalfeeding systems currently employed in entrained flow gasifiers.

Option 1 would require a major R&D effort due to the very limited experience withpressurized, oxygen-blown gasification. A syngas conditioning concept which convertstars, methane and other hydrocarbons needs to be developed.

Option 2, given the development of a suitable feeding system is successful, should, inmany respects, present the easiest solution. Unlike the alternatives, it does not introduceadditional process steps. Research on a dedicated biomass feeding system using screwfeeders and piston compressors is presented in [251]. This system requires particles nosmaller than 1 mm, but since biomass is more reactive than coal, bigger particle sizes mayindeed be acceptable for gasification.

Option 3 offers the advantage that the pretreated biomass can most likely be gasified in aplant designed for coal. That would allow to develop the CCS technology for coal and laterswitch to (co-)gasifying biomass and thereby introduce flexibility regarding CO2 emissionreductions.

As discussed in sections 2.2.2.4 and 2.2.3.11, torrefaction and HTC destroy the fibrousnature of the biomass and thereby make it more suitable for pneumatic feeding. Regardingthe gasification itself, pulverized HTC biocoal from beech wood showed a gasificationbehaviour similar to Rhenish lignite in a laboratory-scale entrained-flow gasifier. Theseexperiments were conducted at atmospheric pressure and 1000–1400°C [203].

Production of a pumpable oil/char slurry by fast pyrolysis is another pretreatment tech-nology which has been suggested to facilitate the entrained flow gasification of biomass[267].

In summary, this chapter offered a general background on the biomass feedstocks andconversion technologies analyzed in this work. Details on the plant designs of the biomassupgrading plants are given in chapter 4, while the BECCS configurations are discussed inchapter 5.

48

3 Scope, methodology and assumptions

The investigations covered in this work break into two parts. The first is the simulation,analysis and comparison of selected biomass upgrading technologies (chapter 4). Thesecond is the simulation and analysis of selected BECCS plants (chapter 5) using both rawand upgraded biomass as a fuel. The study is comparative and attempts to identify whichprocesses and pathways have most merit and under what circumstances. The primaryinterest is on the resulting levelized product costs, and GHG intensities. The core productis upgraded biomass for the plant designs analyzed in chapter 4 and electricity for theBECCS plants in chapter 5.

The selected technologies and configurations are simulated using the software packageAspen Plus. This chemical engineering process simulator offers an extensive componentlibrary of unit operations such as chemical reactors, heat exchangers, pumps and separationprocesses. These serve as building blocks used to compose a process flowsheet with thehelp of a graphical user interface. Aspen Plus then solves the steady-state material andenergy balances and reports the composition, pressure, temperature, specific enthalpy andentropy of all the material streams, and any work and heat flows. Details of the simulationmodels and underlying assumptions are given in section 3.3.

The results from the process simulation allow one to size the plant equipment and estimatethe investment costs and the levelized product costs of upgraded biofuel or, in case ofthe BECCS plants, electricity. Methods, assumptions and data sources for the economicassessment are described in section 3.4.

Exergy analysis is employed to identify and evaluate the thermodynamic inefficiencieswithin the processes and to help uncover potentials for improvement. In addition, exergoe-conomic analysis is applied to the HTC base case design. These exergy-based methodsare explained in section 3.5.

The analysis of the biofuel upgrading processes includes the supply chain cost and GHGemissions, where the supply chain comprises biomass cultivation, storage and transport(section 3.2). For the BECCS plant, supply chain emissions are not taken into account.The net negative CO2 emissions are calculated based solely on the carbon balance of theBECCS plant and, if present, the upstream upgrading plant.

The different upgrading technologies and plant designs are compared based on their costsand GHG emissions per unit of biofuel, energetic and exergetic efficiencies and carbonyield. The carbon yield is a measure of how much carbon is retained by the upgraded bio-fuel and is important for BECCS applications, because it determines the ratio of biomasscarbon which reaches the CCS plant and is therefore available for capture and storage. Inorder to compare the conversion pathways using solid biofuels to those using biogas, it isassumed that solid biofuel pellets displace bituminous coal in an existing pulverized coal-fired power station (PC), while biomethane replaces natural gas in a combined cycle power

49

Chapter 3 Scope, methodology and assumptions

plant (CCGT). In this way, the product of all considered conversion chains is electricity,and the efficiency, avoided GHG emissions and costs of each option can be legitimatelycompared.

The analysis of BECCS centres on gasification. IGCC with fluidized bed gasification ofraw biomass is compared to the entrained flow gasification of torrefied wood and HTCbiocoal. The conversion chains from raw biomass to electricity are analyzed in respect ofelectrical efficiency, carbon capture rate and cost of electricity (COE). For comparison,black box models of PC and CCGT power plants with and without CCS fired on biofuelsand fossil fuels are included. The efficiency, carbon capture rate and specific investmentcost for these plants are taken from literature, and the COE is calculated with the sameeconomic boundary conditions as for the IGCC plants.

The next section describes the scenarios and simulation cases analyzed.

3.1 Scenarios and simulation cases

The term simulation case refers to a plant configuration and its respective Aspen Plussimulation. Data related to material, energy and exergy balances is obtained as a resultof the simulation. A scenario contains additional qualifiers which are relevant for theeconomic analysis and GHG balance. These additional qualifiers include assumptions onthe source of the feedstock, the supply chain, and the plant capacity. Several scenarioscan be derived from the same simulation case.

Since the choice of feedstock can have a large impact on economic viability and GHGbalance of bioenergy processes, a base design for an HTC plant is analyzed for wood,park and gardening waste (PGW), source separated municipal organic waste (MOW),grass and empty fruit bunches from palm oil production (EFB). Due to the differentcompositions and yields, there are dedicated simulation cases for the various feedstocks.For the simulation case number, the first digit identifies the type of feedstock: 1 is wood,2 is PGW with a water content of 50%, 3 is PGW or grass with a water content of 70%,4 is MOW and 5 is EFB.

HTC from park and gardening waste, which seems to be a suitable feedstock, is analyzedwith the most detail, comprising 24 simulation cases. In addition to the base case HTC-3.00, 16 parameter studies, using the same flowsheet as the base case but with modifiedoperating parameters such as temperatures, pressures and reaction times, and 7 simulationcases employing alternative flowsheet designs are investigated.

Wood pelletizing (WP) and torrefaction (TOR) are explored as competing upgrading tech-nologies for wood. Anaerobic digestion with upgrading to biomethane (ADM) and withthe combined production of press cake pellets (ADP) are investigated as two alternativetreatments for grass.

Plant capacity has a significant impact on the product cost due to economy-of-scale effectsfor both the investment and labour costs. Plant capacity is also an important factor forthe supply chain cost and GHG emissions, because larger plants require larger catchmentareas to cover their feedstock demand. To assess the impact of plant capacity, the economicanalysis and the supply chain analysis are conducted for three plant sizes: 2 t/h, 10 t/h

50


Table 3.1: Scenarios for biomass upgrading analyzed in chapter 4. The columns representfeedstocks and the rows represent the base technologies. The cells indicate the scenariosanalyzed for each technology in combination with each feedstock. Scenarios printed inbold are later considered in the conversion chain analysis comprising upgrading andcombustion in a power station (section 4.6). The naming conventions for the scenariosare explained in the text.

wood, SR wood, FR waste grass EFB

HTC HTC-1.0-s/m HTC-3.00-s/mHTC-3.01...3.90-sHTC-4.00-s/mHTC-2.00-s/m

CHPB-3.1...3.3

HTC-3.00-s-G HTC-5.00-s/m*

TOR TOR-1.0-s/m/l

TOR-1-s/m/l

WP WP-1.0-s/m/lWP-1.1-s/m

WP-1.2-m/l

WP-1.3-m/l

WP-1.0-s/m/lWP-1.1-s/m

WP-1.2-m/l

WP-1.3-m/l

ADM ADM-3.0-sADM-3.1-s

ADP ADP-3.0-s

and 50 t/h of feedstock dry matter. In the scenario names, they are denoted as s, m andl, respectively.1 The efficiency is assumed to be identical for all plant sizes, based on theAspen Plus simulations with an input of 2 t/h feedstock dry matter. This is clearly aconservative assumption, as larger plants are likely to have reduced specific heat losses,and the efficiencies of some plant equipment, especially turbomachinery, increases withsize.

The analyzed biomass upgrading scenarios are listed in Table 3.1. Scenarios printed in boldare later considered in the conversion chain analysis comprising upgrading and combustionin a power station (section 4.6). An overview of the analyzed conversion pathways wherebiofuels replace fossil fuels in existing power stations is shown in Figure 3.1.

The scenario names are composed of a multi-letter acronym for the technology, a numberfor the simulation case and an s, m or l to indicate the plant capacity. While mostacronyms are relatively self-evident, CHPB stand for combined heat, power and biocoalproduction and constitutes a wood-fired CHP plant combined with an HTC process. Thefirst digit of the simulation case number indicates the type of feedstock, as explainedabove. The following digits specify modifications to the flowsheet design and/or operatingparameters. 0 or 00 indicates the base case model for the respective technology. Twosources of wood, short rotation (SR) and forest residues (FR), are considered for woodpelletizing and torrefaction. They differ in respect of feedstock cost and supply chain and

1In consideration of the limited feedstock availability, 50 t/h plants are not considered for waste, grass andEFB. To match the maximum available EFB at a typical palm oil mill of approximately 90 ktFM/a, themedium size HTC plant processing EFB has a capacity of 4.5 t/h feedstock dry matter. It is denotedas m*.

51


Table 3.2: BECCS scenarios and their reference cases analyzed in chapter 5. The columnsrepresent the raw feedstocks and the rows represent the power plant types. The cellsindicate the upgrading scenarios analyzed for the respective power plant type in com-bination with each feedstock.

wood, FR wood, SR grass fossil fuel

EF-IGCC with CCS TOR-1.0-l TOR-1.0-m

HTC-1.00-m

bituminous

coal

FB-IGCC with CCS WP-1.0-l raw wood

WP-1.0-m

FB-IGCC without

CCSraw wood

PC without CCS * TOR-1.0-l bituminous

coal

CCGT with CCS * ADM-3.0-s

ADM-3.1-s

natural gas

CCGT without CCS* ADM-3.0-s

ADM-3.1-s

natural gas

* black box power plant model (efficiency and specific investment cost from literature)

are denoted by SR and FR at the end of the scenario names. Grass is assumed to have thesame composition as PGW but differs in respect of feedstock cost and supply chain. ForHTC, the scenario with grass is denoted by a G at the end of the scenario name. ADMand ADP are only run on grass.

With these naming conventions, the simulation case HTC-3.00 is the HTC base designprocessing either PGW-70 or grass. The scenario HTC-3.00-s-G is the small HTC plantbased on this simulation case, processing grass. The scenarion HTC-3.00-m is the mediumsize HTC plant processing PGW-70, based on the same simulation case HTC-3.00. Sincethe scenario qualifiers for plant capacity, type of wood and the distinction between PGWand grass are only relevant for the economic assessment and the supply chain cost andGHG emissions, they are only used in this context. Energy and carbon balance and exergyanalysis are conducted based on the simulation cases.

Apart from the aforementioned combustion of biocoal in a PC power station, the conver-sion chain efficiency is also analyzed for HTC followed by combustion in a medium-scaleCHP plant, and compared to the combustion of raw biomass. The respective simulationcases (not listed in Table 3.1) of the CHP plant are described in section 4.5.15.

BECCS simulation studies and cost estimates are conducted for IGCC plants with en-trained flow (EF) and fluidized bed (FB) gasification. The upgrading scenarios consideredfor the conversion chain analysis are listed in Table 3.2. For comparison purposes, othertypes of power plant with and without carbon capture are included as black boxes. Theseare indicated with an asterisk. An overview of the analyzed BECCS conversion pathwaysis shown in Figure 3.2.

The energy, exergy and carbon balances of the gasification and syngas conditioning sectionsof the IGCC plants are analyzed for some additional cases, comprising modified plantconfigurations and operating parameters (see section 5.1).

52


wood, FR

biomass

bituminouscoal

naturalgas

PGW

grass

wood, SR

EFB

MOW

WP

medium-scaleCHP

PC

CCGT

HTC

ADP

ADM

upgrading power plantdisplacedfossil fuel

CHPB

TOR

no treatment

Figure 3.1: Conversion pathways in which biofuels replace fossil fuels in existing powerstations. A grey background denotes a simplified black box plant model. Dotted linesindicate displacement of fossil fuels.

wood, FR

biomass

bituminouscoal

naturalgas

wood, SRWP

ADM

upgrading power plant fossil fuel

TOR

HTC

no treatment

FB-IGCC-CCS

EF-IGCC-CCS

FB-IGCC-noCCS

grass

CCGT-CCS

CCGT-noCCS

PC-noCCS

Figure 3.2: Conversion pathways with BECCS and their associated reference scenarioswith fossil fuel use and biomass without carbon capture. A grey background denotes asimplified black box plant model.

53


3.2 Supply chain cost and GHG emissions of biofuel production

The supply chain for biofuel production and utilization in an existing power plant is shownin Figure 3.3. The green blocks are the operations, comprising cultivation and harvest,storage, biomass upgrading, combustion in the power plant, and required transport ofbiomass and biofuel. GHG emissions and costs are calculated for each step. The redblocks denote the biomass or biofuel at various stages in the process.2 The blocks on theleft designate input parameters, where purple indicates GHG emissions data, blue is costdata, and grey is reserved for miscellaneous data and assumptions required for the models.

CH4 and N2O emissions are accounted for using their 100-year global warming potential(GWP100) of 25 and 298 CO2 equivalents, respectively [268].

A preliminary analysis of HTC from short rotation wood grown in Germany showed thatthe biocoal costs are largely determined by the feedstock cost. Therefore a second setof scenarios is analyzed for upgrading wood, where the upgrading plant is situated in aNorth American location where cheap forest residues are available in large quantities. Theresulting biofuel is then shipped to Europe. Since the aim is to investigate the trade-offbetween cheap feedstock and the additional long distance transport of the biofuel, allother economic parameters remain unchanged. Because preliminary calculations showedthat for HTC, waste feedstocks are more viable than wood, the North American scenariosare only analyzed for conventional wood pelletizing and torrefaction, which are both wellsuited to wood.

A more specific case study is conducted for HTC from palm oil production waste. Inthis case, the HTC plant is situated in Malaysia, and economic parameters such as plantoperators wages and electricity price are duly adjusted.

The supply chain operations are explained in detail in the following sections.

3.2.1 Biomass cultivation and harvest

The “cultivation and harvest” block in Figure 3.3 comprises all operations required toprovide the biomass at the farm gate.3 Cost and GHG emissions during cultivation andharvest may occur through the use of fertilizers and pesticides, water for irrigation, ma-chinery and labour and the purchase of arable land and seeds. GHG emissions may alsoresult from changes in the soil carbon balance due to land use change.

Assumptions on composition, heating value, cost and GHG emissions from the cultivationand harvest of the feedstocks are given in Table 3.3. The cost of poplar chips from shortrotation forestry (SR) in Germany range from 67–82 €/tDM [269], 76.6 €/tDM is assumedis this work. Cost for wood chips from forest residues in North America are based on

2The term biomass, or feedstock, denotes the raw biomass, while the term biofuel denotes the upgradedbiomass after pretreatment. The latter comprises HTC biocoal pellets, torrefied wood pellets, conven-tional wood pellets, and biomethane.

3The term “farm gate” is used here to refer to the point of hand-over where the biomass is picked up fortransport to the upgrading plant. The same term is also used more generically to cover the point ofhand-over from forestry operations or waste collection. Note that for wood and grass, the transportfrom field to farm gate for reason of simplicity is included in the block “transport to the biofuel plant”.

54


biomass cultivation and harvest

biomass storage

truck transport to harbour

truck transport to power plant

ship transport

GHG emissions [t /t ]CO2,eq biomass

dry matter loss

dry matter loss

GHG [t /km/t ]CO2,eq biofuel



GHG [t /km/t ]CO2,eq biomass

fossil fuel GHG [t /t ]CO2,eq fossil fuel

investment cost data

electricity upstream GHG [t /MWh ]CO2,eq el

process related CH emissions 4

cost [€/t ]biomass

cost [€/t ]biomass

cost [€/km/t ]biofuel



distance

distance

distance

yield [t /ha]biomass

electrical efficiency

modelling assumptionsexperimental data

cost [€/km/t ]biomass

fossil fuel cost [€/km/t ]fossil fuel

labour requirement + cost electricity cost

transport to biofuel plant

power plant

upgrading plant

biomass at farm gate

biomass at upgrading plant gate

biofuel at upgrading plant gate

biofuel at harbour (Europe)

biofuel at power plant gate

avoided GHG

transport distance

biofuel yield, auxiliary energy demand

in Europe overseas

equipment sizes

radius model

Aspen Plus simulation

plant cost estimate

substitution of fossil fuel by biofuel

Figure 3.3: Schematic of the biofuel supply chain.

55


values reported for Georgia [270]. The composition for both types of wood is assumed tobe identical is this study4, although in reality it varies slightly depending on the species.

Three types of waste are considered: park and gardening waste with a water content of70% (PGW-70), park and gardening waste with a water content of 50% (PGW-50), andMOW. MOW differs from park and gardening waste by its higher ash content.5

The costs of waste are negative, because a remuneration is paid for treatment. As a benchmark, treatment cost for composting lie in the range of 30–80 €/tFM for MOW and 15–30€/tFM for park and gardening waste [271, page 38].

Grass is assumed to have the same composition as PGW-70. The cost are estimated usingdata from [272] for extensively cultivated permanent grassland in Germany with a yieldof 5.8 tDM/ha/a.6

The measured composition of the EFB used for HTC experiments by Stemann [23] showsa high hydrogen content compared to other types of biomass, which leads to a highercalorific value. The cost of EFB is difficult to estimate because there is no establishedmarket. Quotes for charges of 100–1000 tonnes on an internet trading platform7 suggest4–14 USD/tFM. Shuit [67] reports 2.8 USD/tFM for 2005 (citing [274]). Another way toassess the cost of the EFB is the opportunity cost, i.e. their value when used as a soilconditioner on the oil palm plantation. According to Ravi Menon et al., [64], the netreturn from EFB used for mulching is 4.45 €/tFM (14.4 RM/t) taking into account thefertilizer value, the cost associated with its application and the increased palm oil yield.A cost of 4.00 €/tFM is assumed in this work.

The embedded GHG emissions for growing SR poplar range between 4–55 gCO2,eq/GJdepending on whether they are grown on acre land or grassland, and the extent thatdirect and indirect land use change is taken into account [275]. Preliminary calculationsshowed that this is equivalent to 0.1–1% of the total supply chain emissions of HTCor torrefaction. Given their minor contribution, the embedded emissions of SR woodcultivation are neglected in this work.

For grass from extensive permanent grassland, it is assumed that the grass is fertilisedwith digestate only [56], and that the embedded GHG emissions are negligible.

For waste, the GHG emissions depend on the reference scenario, i.e. what would happenwith the waste were it not used as a feedstock for upgrading. The GHG emissions ofthe reference scenario are accounted for as avoided emissions in the supply chain of theupgrading plants. For PGW and MOW, the reference scenario is composting. It is assumedthat GHG emissions from composting are negligible, since the embedded emissions fromthe energy consumption at the composting works and the positive effects due to humusformation and nutrient replacement compensate each other [276].

4Based on the measured composition of poplar used in HTC experiments conducted by Stemann [185].5Ash and moisture content of MOW are based on [33]. No data was found in literature for the dry

and ash free composition of waste, therefore it was estimated, with an O/C ratio higher than for woodbecause of the lower lignin fraction.

6Land subsidies are obtained for the cultivation of extensive grassland because of its importance forbiodiversity. Premiums depend on location, agricultural practices and biodiversity value [273, page 62].It is assumed here that land cost (tenure) and subsidies compensate each other.

7http://www.alibaba.com, accessed 01-Sep-2011 and 07-Jan-2013.

56


Table 3.3: Biomass feedstocks used in this work.

wood

chips

SR

wood

chips

FR

PGW-50 PGW-70 MOW grass EFB

water (w.b.) [w%] 50% 50% 50% 70% 70% 70% 65%

carbon (d.b.) [w%] 47.49% 47.49% 42.54% 42.54% 36.05% 42.54% 49.18%

hydrogen (d.b.) [w%] 6.42% 6.42% 5.94% 5.94% 5.03% 5.94% 7.00%

nitrogen (d.b.) [w%] 0.07% 0.07% 1.14% 1.14% 0.97% 1.14% 0.64%

sulphur (d.b.) [w%] 0.05% 0.05% 0.16% 0.16% 0.14% 0.16% 0.07%

oxygen (d.b.) [w%] 44.18% 44.18% 44.62% 44.62% 37.81% 44.62% 38.59%

ash (d.b.) [w%] 1.78% 1.78% 5.60% 5.60% 20.00% 5.60% 4.52%

HHV (d.b.) [MJ/kg] 19.548 19.548 17.115 17.115 14.183 17.115 21.332

LHV (w.b.) [MJ/kg] 7.855 7.855 6.690 3.038 2.218 3.038 5.350

bulk density

(w.b.)

[kg/m3] 290 290 100 167 750 167 350

cost (farm gate) [€/tFM] 38.31 16.54 -20.00 -20.00 -50.00 22.45 3.00

cost (farm gate) [€/GJHHV] 3.92 1.69 -2.34 -3.90 -11.75 4.37 0.40

GHG emissions

(farm gate)

[kgCO2,eq/tDM] 0 0 0 0 0 0 -3297

For EFB from palm oil production, the reference scenario is assumed to be dumping. Themethane emissions were calculated with the IPCC default method (tier 1) [277] assuminga methane correction factor of 80% for solid waste disposal sites with a depth of morethan 5 m, a fraction of 50% methane in the generated gas, and a fraction of dissimilateddegradable organic carbon of 50% [278]. 0.5% of the EFB nitrogen is assumed to beconverted to N2O [37]. The resulting GHG emissions amount to 3297 kgCO2,eq/tDM ofEFB. GHG emissions related to growing the oil palm are not included, because they shouldbe assigned to the palm oil rather than the EFB.

3.2.2 Storage

The availability of biomass crops is seasonal, but the upgrading plants are assumed tooperate year round. Therefore, storage over several months is required. Storage is charac-terized by the dry matter loss and the cost for land, construction of the storage facilities,and labour and machinery to load and unload the material into and out off the storagefacility.

The dry matter loss by biological decomposition depends on the type of biomass andits moisture content and particle size. For fresh wood chips, dry matter loss throughbiological decay can exceed 20% for a storage period of 7–9 months. For logs, the lossis usually limited to 1–3% per year [28, pages 290–291]. Due to the smaller particle sizefrom harvesting, the loss is higher for short rotation wood than for forest residues [76].

Some drying also takes place during storage, especially in the summer months, but this isnot taken into account in this work.8

8For processes that require drying, the annual energy demand may therefore be overestimated.

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Table 3.4: Transportation distances in the biomass supply chain.

product from to vehicle distance [km]biomass field farm gate agricultural vehicle variesbiomass farm gate upgrading plant truck or agricultural vehicle variesbiofuel upgrading plant harbour, Malaysia truck 100biofuel upgrading plant harbour, America truck 200biofuel harbour, Malaysia harbour, Europe ship 16000biofuel harbour, America harbour, Europe ship 7000biofuel harbour, Europe power plant truck 100biofuel upgrading plant power plant truck 100biomethane upgrading plant power plant pipeline 100

Since HTC takes place in water and a high feedstock moisture is not an issue, ensilingthe wood chips presents an interesting option. However, Wirth et al. conclude that thehigh costs for the foils and the construction of silos outweigh the savings from reduced drymatter loss [76].

Grass is assumed to be stored as silage.

Since waste is generated year round, it is assumed that long-term storage is not required,and dry matter losses and storage costs are neglected.

The assumptions regarding storage loss and cost are given in section A.1.1.

3.2.3 Transport

The supply chain model includes transport of the biomass to the upgrading plant andtransport of the biofuel from the upgrading plant to the power plant. The various transportoperations are illustrated in Figure 3.3. The transport costs are calculated based ondistance, and the charter and fuel cost of the respective vehicles. Details are provided insection A.1.2. Assumptions on the respective transport distances are given in Table 3.4.The calculation of the transport distance of the raw biomass to the upgrading plant isexplained in the next section. Transport cost and GHG emissions for different types ofvehicles employed in the biomass supply chains are given in Table A.3 and Table A.6,respectively.

3.2.3.1 Transport of the biomass to the upgrading plant

The cost for on-field transport to the farm gate is estimated based on the field size. Detailsare provided in section A.1.2.2.

The transport distance from farm gate to the upgrading plant is dependent on the availab-ility of the feedstock in the surroundings and on the plant capacity. Larger plants requirelarger catchments. The required catchment area for crops (SR wood and grass) is calcu-lated from the yield per hectare of the respective crop and an assumption on the share ofthe surrounding land which provides feedstock for the upgrading plant (f1). The average

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Table 3.5: Assumptions for the calculation of the transport distance to the upgradingplant processing wood or grass, and the resulting required catchment area.

wood SR wood FR grassyield [tDM/ha/a] 10 2.5 5.8dry matter loss, harvest [% p.a.] 3% 3% 8%share of land supplying feedstock (f1) [–] 10% 50% 10%

transport distance is calculated based on Kappler [279, page 61], assuming that the up-grading plant is situated in the middle of its circular catchment area, and that croplandsupplying the plant is evenly distributed. A factor for the deviation between actual andradial transport distance is taken into account. Further details on the calculation modelare provided in Table A.1.2.1 and [76].

Transport from the farm gate to the upgrading plant is conducted by an agriculturaltruck/trailer vehicle or truck, whichever is cheaper (see section A.1.2.2). Transport bytruck requires reloading from the agricultural vehicle used for the transport on the field,and thus is only worthwhile for transport distances above 13 km for wood chips [76].

Assumptions on yields, dry matter loss during harvest and the share of available land inthe surroundings are summarized in Table 3.5.

Waste biomass is collected from households and parks and brought to a disposal facility,irrespective of the treatment technology. The cost of collection is not taken into accountin this work. However, due to the larger capacity of an HTC plant compared to a typicalcomposting plant, longer transport distances for the waste may occur. It is assumedthat no additional transport cost is incured when the waste processing capacity of theupgrading plant does not exceed that of a typical composting plant, namely 10 kt/a forplants processing only park and gardening waste, and 20 kt/a for plants processing MOWas well [280]. If the capacity of the upgrading plant is higher, the additional cost fortransporting the waste over longer distances is included in the biofuel production cost.This transport distance is calculated based on the population density, the amount ofwaste generated per person, and a factor describing how much of the generated waste canbe acquired. Details are given in section A.1.2.3.

No transport cost are assumed for EFB, since the HTC plant is assumed to be situatednext to the palm oil mill. One mill provides approximately 20–90 kt/a of EFB (seesection 2.1.2.5). If other mills are close by, additional EFB could be acquired to increasethe capacity of the upgrading plant, but this option is not considered in this work.

3.2.3.2 Transport of the biofuel to the power plant

Domestically produced biofuel pellets are transported to the power station by truck. Bio-fuels produced overseas are transported to a harbour by truck, shipped to Europe, and thentrucked to the power station. The respective transport distances are given in Table 3.4.

The costs for the biofuel transport by ship from Malaysia and North America are estimatedbased on the distance, charter cost for the vessel, fuel consumption, freight capacity, feesfor harbours and canals and loading/unloading cost. Details are given in section A.1.2.4.

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Table 3.6: Cost of solid biofuel shipment by bulk cargo to a European harbour.

origin cargo [kt] distance [km] cost [€/t]HTC pellets Malaysia 29.750 16.000 35.79HTC pellets, torrefied pellets North America 63.000 7.000 17.56wood pellets North America 63.000 7.000 18.63

The results, summarized in Table 3.6, are in good agreement with values reported inliterature. For example, Obernberger and Thek report shipment costs of 31.7 €/t forpellet transport from Indonesia to Italy [281]. Sikkema et al. report shipping costs forwood pellets from north America to Rotterdam of below US$25 for long-term contractssigned in 2009 [10].

Shipping prices are highly volatile due to the supply and demand of shipping capacity.Between 2004 and 2008, the charter rates for Capesize vessels rose fivefold due to a shortagein shipping capacity [104]. Shipping costs are also strongly dependent on the size of ships,whereas the maximum ship size is limited by harbours and canals. Another importantfactor is whether a well established trade route is used or not [104]. Between 2002 and2010, freight costs for transporting pellets from North America to Europe ranged from 27€/t to 69 €/t [10]. Assumptions on future shipping prices are therefore subject to a highdegree of uncertainty.

The cost for the pipeline transport of biomethane to a power plant is assumed to be 1.11€/GJHHV based on [282].

3.2.4 Processing at the biofuel plant

The upgrading plants consume electricity as an auxiliary fuel. Embedded emissions forelectricity are based on the power mix in Germany and Malaysia (see Table A.6). Forupgrading plants that co-produce a surplus of electricity, a GHG credit is taken intoaccount based on the avoided generation in the national power mix.

Emissions embedded in the plant equipment are neglected.

Process emissions comprise methane emissions from the anaerobic digestion and emissionsrelated to the spreading of the digestate on fields (see section 4.4.6).

It is assumed that any carbon compounds in the waste water are completely oxidizedduring treatment, thus there are no methane emissions to be considered.

3.2.5 Combustion of the biofuel at a power plant

Solid biofuels are assumed to displace bituminous coal in a pulverized coal-fired powerstation. Due to the lower calorific value of the biofuels compared to bituminous coal,the efficiency of the combustion is slightly lower. Based on simulation models of thecombustion9 of bituminous coal, biocoal, torrefied wood and wood pellets the efficiency of

9Assuming an air ratio of 1.15, an unburned carbon fraction of 0.3%, a radiative heat loss of 0.5%, andan exhaust gas exit temperature of 130°C.

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exhaust gasair

power plantupgrading plant

air

biomass

combustionupgrading

process

steam cycle

biofuel

materiallosses

ash

WSC

QSC

WUPG

Figure 3.4: Schematic for the efficiency definitions of biocoal utilization in a power plant.

the combustion ηcomb is calculated:

ηcomb =QSC

mbiofuelHHVbiofuel(3.1)

A schematic of the relevant material and energy flows is shown in Figure 3.4. With asteam cycle efficiency ηSC of 42.5%10 and the respective energy yields and auxiliary energyconsumptions of the upgrading process WUPG, the efficiency of the power plant ηPP andof the overall conversion chain ηCC from raw biomass to electricity is calculated:

ηSC =WSC

QSC(3.2)

ηP P = ηcombηSC (3.3)

ηCC =WSC − WUP G

mbiomassHHVbiomass(3.4)

The avoided GHG emissions per MJ biofuel Δbbiofuel can be calculated as

Δbbiofuel =ηbiofuel

ηfossilbfossil − bbiofuel (3.5)

where bbiofuel and bfossil are the specific emissions per MJ related to the supply and com-bustion of biofuel and fossil fuel, respectively. bbiofuel comprises the upgrading process,transport of the biomass and biofuel, and harvest of the biomass. bfossil comprises theCO2 emissions from the combustion and the supply chain GHG emissions based on datafrom literature. The emission factors of bituminous coal, pulverized lignite and naturalgas are given in Table A.6.

The GHG mitigation cost cGHG is defined as the additional cost for using biofuel instead

10This results in an energetic efficiency for the bituminous coal-fired power plant of 40% (LHV) or 38.3%(HHV).

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Table 3.7: Assumptions on fossil fuel prices.

bituminous coal [€/GJHHV] 2.52pulverized lignite (delivered milled and dried) [€/GJHHV] 4.70natural gas [€/GJHHV] 6.31

of fossil fuel divided by the avoided GHG emissions:

cGHG =cbiofuel − ηbiofuel

ηfossilcfossil

Δbbiofuel(3.6)

where cbiofuel and cfossil are the specific cost per MJ of biofuel and fossil fuel at the powerplant gate.11 The cost of bituminous coal, pulverized lignite and natural gas are given inTable 3.7.

Although pelletized biofuels are much easier to handle than raw biomass, their use in anexisting coal-fired power station may require modifications at the fuel preparation andfeeding systems. The extend of the modifications will likely depend on the quality of thebiofuel and the co-firing rate. The cost for such modifications are not considered in thiswork.

3.3 Process simulation

The biomass upgrading plants and the IGCC plants except for the steam cycle are mod-elled with the process simulator Aspen Plus V7.1. The steam cycles of the IGCC plantsare modelled with the simulation packages GateCycle Version 5.52.0r and EBSILON Pro-fessional 9.00.12

3.3.1 Property methods

Chemical process simulations are heavily reliant on the quality and suitability of thethermophysical property methods used. Selecting an appropiate method for the givenchemical system is therefore essential.

Key material subsystems in the plant models include the biomass and biocoal slurries inthe HTC plant, hot gas mixtures in the gas path of the IGCC plant, and pure water/steamin the steam cycles. Different property methods are used to model these respective fluidsand the solid feedstocks and biofuels.

3.3.1.1 General thermophysical property methods for the simulations

Liquid flow streams and recovered flash steam in the HTC process are composed mostlyof water (>95 mol%), but contain small amounts of acetic and formic acid and dissolved11MOW contains 20% ash, therefore high ash disposal costs may have a negative influence on the economic

performance. The cost of the ash disposal at the power plant is included in cbiofuel.12GateCycle is used for the simulation cases EF-IGCC-HTC-1, EF-IGCC-TOR-1, EF-IGCC-WP-1. EB-

SILON Professional is used for FB-IGCC-wood-1, FB-IGCC-wood-0 and FB-IGCC-WP-1

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CO2. Therefore, steam table functions are not applicable. The NRTL-HOC propertymethod is used for the HTC process, since the Hayden-O’Connell (HOC) equation ofstate is recommended for mixtures containing carboxylic acids such as acetic acid [283].Solubility of CO2 is calculated using Henry’s law, which is accurate for acid solutions withpH-values below 5 [284]. The pH value of HTC process water was found to be 3.4–4.5 inlaboratory experiments [163]. The temperature dependent Henry constant is calculatedwith coefficients from the Aspen Plus APV71 ENRTL-RK data base, since this model isvalid for temperatures up to 226°C.

The NRTL-HOC property method is also used for the other biomass upgrading processes.

For the IGCC plants modelled with Aspen Plus, the property method RKS-BM (Redlich-Kwong-Soave with Boston Matthias Alpha function) is used for the gas path, while waterand steam are modelled with the steam table function STEAMNBS. In GateCycle andEBSILON Professional, the combustion gas, being near ambient pressure, is modelled asan ideal gas, and water/steam is modelled with the steam table formulation IAPWS-IF97.

For the calculation of chemical exergy, the thermodynamic environment described bySzargut [285] is employed.

3.3.1.2 Validation of the thermophysical property method for the HTC simulation

For the design and analysis of the heat recovery scheme in the HTC process, it is importantthat the properties of liquid and gaseous water in the temperature range of 100–220°Cand close to saturation pressure are modelled realistically. Since the NRTL-HOC methodis generally not recommended for pressures exceeding 10–15 bar [283], and pressures ofup to 35 bar occur in the HTC plant models, the properties of water in the relevanttemperature range of 100–220°C are compared to those resulting from the the steam tablefunction STNBS. The difference amounts to −3 to +46 kJ/kg for the specific enthalpy ofboiling water, −33 to +66 J/kg/K for the specific entropy of boiling water, and +1 to +8kJ/kg for the enthalpy of evaporation. For liquid water at 35 bar, the difference is 0 to+47 kJ/kg.

To assess the effect of the property method on the HTC plant efficiency, a simulation ofa preliminary HTC plant model was conducted without gaseous and liquid byproductsand run first with NRTL-HOC for the overall plant and second with STEAMNBS (andideal gas for the gas path of the boiler and drier). Although the mass flows of recoveredflash steam differed by up to 6%, the difference in the overall efficiency of the HTC plantamounts to only 0.03 percentage points and is therefore negligible.

The extent of CO2 solubility in the liquid products of the HTC reactor is important,because the amount of gaseous CO2 largely determines the amount of steam lost withthe gaseous byproducts, and this represents a significant energy loss from the reactor (seesection 4.5.3). The model results were therefore compared to data from literature. At200°C and 25 bar, 23% of the CO2 is dissolved in the Aspen Plus model, while with datafrom Butler [284], 29% is dissolved.

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3.3.1.3 Characterization of feedstocks and solid biofuels

The dry matter fractions of biomass and solid biofuels are modelled in Aspen Plus asso-called nonconventional solids, characterized by their elemental composition (C, H, O,N, S, ash), higher heating value (HHV) and heat capacity. The moisture is modelled asliquid water.

The HHV is calculated based on the elemental composition with a correlation suggestedby Channiwala and Parikh [286]. The correlation is reported to be valid for a wide rangeof compositions and fuels, including coals, biomass and char as well as liquid fuels, withan average error of 1.45%.

A temperature dependent function for the heat capacity c of the form

cs,DM = aT + b (3.7)

for the dry matter cs,DM is used for biomass, biocoal and ash. The parameters a and bfor biomass are derived from data for wood [287, 288]. Biocoal is assumed to have a heatcapacity similar to that of coal, and the parameters a and b for biocoal and ash are basedon data for fossil coal [289].

The heat capacity of the moisture is modelled as liquid water, cp,H2O,l, so that the heatcapacity cs,FM of the wet feedstock reduces to

cs,F M = cs,DM · (1 − w) + cp,H2O,l · w (3.8)

In reality, the moisture bound in the biomass or biofuel will have different propertiesfrom liquid water. For coal, the error made by assuming that the heat capacity of themoisture equals that of liquid water was found to be negligible for practical purposes [289].For biomass, the employed heat capacity function is based on the apparent specific heatcapacity of the dry matter at full hydration [287]. This is a good description of the rawbiomass and the slurry in the HTC plant, which have a water content well above fullhydration.13

The physical exergy of the dry matter is calculated using the heat capacity function.The chemical exergy is calculated based on the elemental composition and HHV with acorrelation from [290].

More details on the properties of solid biomass and biofuels are given in section A.2.1.1.

3.3.2 Models for process units and general assumptions

In this section, the modelling assumptions for the process units are explained. The ra-tionale behind the flowsheet designs and modelling assumptions specific to the individualplant designs are discussed in chapters 4 and 5.

13Wood is fully hydrated at a water content of approximately 20%.

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3.3.2.1 General assumptions

Ambient air is assumed to be at 15°C, 1.013 bar, and 60% relative humidity.Modelling assumptions on standard process equipment such as heat exchangers and tur-bomachinery, and the heat and pressure losses of the components are given in sectionsA.2.1.2 to A.2.1.4.

3.3.2.2 Screening

Screening is required when the biomass contains impurities such as bits of metal or plastic.This is especially indicated for waste streams like MOW. The energy consumption ofthe screening processes is neglected in the analysis, but equipment and labour costs foroperating personnel are taken into account for biomass upgrading plants which processPGW and MOW. The screening is assumed to comprise a shredder, drum screen, magneticseparator and air classifier.

3.3.2.3 Size reduction

The specific electricity demand for milling depends on the feedstocks as well as the requiredparticle size and the type of mill. In the simulation models, the particle size distributionis not modelled, and milling is simply characterized by its electricity demand based ondata from literature. Details are given in section A.2.1.5.

3.3.2.4 Pelletization

Electricity consumption of the pellet press is assumed to be 180 kJ/kg based on vendorsinformation for wood pellet presses reported in [143].

3.3.2.5 Biomass pressurization

Biomass is mixed with water to create a pumpable slurry in the HTC plant base design.In some of the alternative HTC designs analyzed in section 4.5.12.2, a plug-forming feederis employed, which can pressurize the raw biomass as delivered with a dry matter con-tent as high as 30%, without adding water. In the simulation model, the pressurizationequipment is characterized by the maximum allowed solid content of the slurry and thepump efficiency or specific electricity consumption. Background information on the choiceof slurry pressurization systems is given in section 4.5.2.In the IGCC plant designs, dry biomass is pressurized via lock hoppers or piston feedersunder an inert atmosphere. Nitrogen is used as inert gas for pressurization. These systemsare characterized in the simulation model by their inert gas consumption. The electricitydemand results from the compression of the inert gas to the required pressure of 60 barin the simulation.Lock hoppers can operate at pressures of up to 30–40 bar [249, page 197], thereby limitingthe maximum operating pressure of the gasifier.Details of the modelling assumptions are given in section A.2.1.6.

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3.3.2.6 Mechanical dewatering

Bulk water can, to some extent, be removed from the feedstock in liquid form by mech-anical pressure. Mechanical dewatering consumes much less energy than thermal drying,where the feedstock moisture has to be evaporated. The extent to which mechanicaldewatering is possible therefore depends on the form in which water is present in thefeedstock. In hygroscopic materials such as raw biomass, peat and lignite, the moisture ismostly bound by capillary absorption or enclosed in cell walls, and can only be removed bythermal drying. Hydrothermal treatment releases this water, thereby enabling mechanicaldewatering to a greater extent. HTC biocoal was successfully dewatered to a dry mattercontent of 55–68% with a laboratory press (see section 2.2.3.11).

The dewatering technologies deployed in the simulation models in this work comprise afilter press for dewatering the HTC biocoal, a screw press for separating the solid presscake and the liquid phase in the case of hydrothermally conditioned grass silage, and adecanter for the fermentation residues from anaerobic digestion. The technologies arecharacterized by the final dry matter content of the solid product, the share of feedstockending up in the liquid phase, and their electricity consumption. The assumptions for thesimulations are based on data from literature. Details are given in Table A.17.

3.3.2.7 Thermal drying

Drying is a complex process involving combined heat and mass transfer, taking place atthe surface and inside of the particles being dried. For this work, a rather simple modelis applied. The amount of drying medium required to evaporate water from the feedstockand give the desired moisture content of the final product is calculated using material andenergy balances under equilibrium conditions.

The different drying processes employed in the plant models are characterized by thedrying medium and operating temperature. They comprise low temperature drying usingair at about 90°C as the drying medium, high temperature drying using combustion gasat 500°C and superheated steam drying (SSD). The systems are described in detail insection 4.1 and the modelling assumptions are given in section A.2.1.8.

The drying model employed does not take into account the temporal profile of the dryingprocess. Thus, differences in the drying behaviour of wood chips and biocoal due to particlesize and moisture transfer mechanisms within the particle are not modelled. Generally, asmaller particle size allows for shorter drying times, and thus a more compact drier design.

The drier model assumes that all water behaves like bulk water. In reality, the propertiesof bound water in hygroscopic materials may significantly deviate from those of bulkwater, because the forces binding the water to the solid must be overcome in the dryingprocess. For lignite, the bonding forces can usually be neglected for drying to a watercontent of approximately 17%. When drying lignite from 20% to 10%, on the otherhand, the additional energy required to overcome the bonding forces is equivalent to 20%of the enthalpy of evaporation [291, page 54]. Although HTC biocoal resembles lignitein terms of its composition, it is not known whether it has a similar drying behaviour.Generally, the drying characteristics have to be determined experimentally, and to theauthor’s knowledge no data has been published on the drying of HTC biocoal. The fact

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that, unlike raw lignite, HTC biocoal can be mechanically dewatered to a high extent,indicates that the moisture is present in different forms. Fortuin reports that for dryingwood, the water-binding enthalpy can be neglected [292, page 47]. Simulation resultsindicate that for drying wood from 28% to 10%, the water-binding enthalpy is equivalentto 4% of the enthalpy of evaporation [292, page 177].

A heat loss of 5% is assumed for all drier models. Carbon losses by evaporation of volatilesare neglected. They are usually lower than 0.1% for drying temperatures below 120°C butcan exceed 1% at higher temperatures [293, page 297].

3.3.2.8 HTC, torrefaction, anaerobic digestion

The HTC, torrefaction and anaerobic digestion reactors are modelled as black boxes withgiven yields based on published experimental data. Details can be found in the respectivechapters. The reactor volumes for the HTC and anaerobic digestion reactors are requiredfor the heat loss calculations and investment cost estimates and are calculated based onthe residence times. The heat loss is calculated using reactor dimensions and the insideand outside temperatures. For details see sections A.2.1.4 and B.3.4. The heat loss of thetorrefaction reactor is assumed to be 3% of the torrefaction gas energy based on HHV.Pressure losses are given in section A.2.1.4.

The electricity demand for the stirrer in the AD reactor is assumed to be 3% of the energyof the produced biogas, based on [28, page 894].

3.3.2.9 Aerobic waste water treatment

The electricity demand for the aerobic treatment of HTC waste water is estimated usingthe oxygen demand for the oxidation of the dissolved organic compounds. It is assumedthat the biological oxygen demand (BOD) equals the stoichiometric oxygen demand forcomplete oxidation. The O2 provided by the aerator is assumed to be twice the BOD. Theaeration efficiency can vary between 0.6–7.3 kgO2/kWhel [294] depending on the aerationsystem. A value of 1.825 kgO2/kWhel is assumed for this work.

3.3.2.10 Boilers and furnaces

Several types of combustion reactor are employed in the plant models. Stoker boilers forthe combustion of upgraded solid biofuels or wood are used to provide saturated steamor superheated steam in the various biomass upgrading plants. For the comparison ofwood and biocoal-fired CHP plants in section 4.5.15, fluidized bed boilers are employed.Pulverized coal firing is presumed for the combustion of upgraded biofuels in existingcoal-fired power stations in section 4.6.

Furnaces fired on upgraded biofuels or raw wood supply hot gas for the drier in the woodpelletizing and torrefaction plants. Offgas incinerators are employed to oxidize carbonmonoxide and volatile organic compounds in the gaseous byproducts from some of theHTC plant designs.

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Combustion is assumed to be complete except for a specified fraction of unburned carbonfrom solid fuels which leaves the reactor with the ash. The gaseous reaction productscomprise CO2, H2O, SO2 and N2. The unburned carbon fraction, air ratio and radiativeheat loss are dependent on the type of boiler and are given in Table A.19.

The minimal exhaust gas outlet temperature for wood boilers without flue gas condens-ation is assumed to be 120°C. Depending on the plant designs, higher exhaust gas tem-peratures may occur because thermal energy demand is confined to higher temperaturelevels. Lower exhaust gas temperatures may occur where the exhaust gas is used for lowtemperature drying or hot water production in boilers with flue gas condensation. Detailsare given in the descriptions of the respective plant designs.

For the incineration of the low calorific offgas leaving the HTC reactor, a minimum com-bustion temperature of 800°C is ensured by regenerative preheating of the reactants.

3.3.2.11 Gasification

Two types of gasifiers are employed in the plant models: an entrained-flow gasifier withan operating temperature of 1550°C and a fluidized-bed gasifier operated at 900°C. Inhigh temperature gasification, the reactions come close to chemical equilibrium. Tremelet al. conducted gasification experiments with HTC biocoal using a bench-scale entrainedflow gasifier at atmospheric pressure, and conclude that, at 1400°C, the gas compositioncan be predicted by chemical equilibrium [203]. Steam-blown entrained-flow gasificationof torrefied wood at 1400°C was also shown to be described satisfactorily by chemicalequilibrium [295]. For the commercial scale Shell coal gasifier operated at 1370°C, all gasspecies were predicted by chemical equilibrium within ± 0.7% of their measured values[296].

The entrained flow gasifier is therefore modelled using chemical equilibrium, except fora specified fraction of unconverted carbon which leaves the reactor with the ash. Thegaseous reaction products comprise H2, CO, CH4, CO2, H2O, H2S and N2.

The gas composition from the fluidized-bed gasification of biomass cannot be satisfactorilypredicted by chemical equilibrium [249, page 10], and approach temperatures need to beintroduced into the equilibrium calculations. The approach temperatures for the individualgasification reactions are calibrated based on published experimental data obtained fromthe IGT/Renugas gasifier under pressurized, oxygen/steam blown operation [297, 298].Details on the model are given in section A.2.1.10, and the results and implications of thismodel are discussed in section 5.1.3. The reaction products are the same as for entrainedflow gasification, plus tar. Tar formation is assumed to be 2.77 g/kg of dry feed. This isequivalent to 2.6 g/m3

STP in the simulation model FB-wood-1.14

Tar is modelled as equal parts of phenol (C6H6O), toluene (C7H8) and pyrene (C16H10 ),to approximate different tar classes (see [258, page 49]).

In both the entrained flow and fluidized bed gasification model, the steam to fuel ratio isset, while the oxygen is adjusted to reach the desired gasification temperature. A minimumsteam to fuel ratio of 0.06 kg/kgfuel is assumed.14The assumed tar formation is based on a tar content of 2–3 g/m3

STP measured in atmosphericsteam/oxygen-blown gasification experiments using in-bed dolomite [299].

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The modelling assumptions for both gasification processes, including heat, pressure andunburned carbon losses, are given in Table A.20.

3.3.2.12 Syngas and biogas cleaning and conditioning

Process units employed in the syngas cleaning and conditioning sections of the IGCCcomprise a methane steam reformer, water gas shift reactor, wet scrubber, hot gas desul-phurization unit, and acid gas removal unit. The flowsheet design is explained in detail inchapter 5.

For anaerobic digestion with biomethane injection into the natural gas grid, CO2 is re-moved from the biogas by pressurized water scrubbing.

The methane steam reformer and water gas shift reactor are modelled as equilibriumreactors, the wet scrubber is modelled as a flash process and the hot gas desulphurizationunit, acid gas removal unit and pressurized water scrubber are modelled as black boxeswith a separation efficiency and energy consumption based on data from literature. Detailsare given in section A.2.1.11.

3.3.2.13 CO2 compression

Compression of the captured CO2 to 110 bar for transport and storage is included in theplant models, based on [300].

3.3.2.14 Gas turbine system

The gas turbine systems employed in the IGCC plant models are fired on a hydrogen-richsyngas. In order to fit the simulation models to published operating data, the gas turbinesystems were first simulated fuelled with natural gas (modelled as CH4).

The performance of gas turbine systems is strongly dependent on their capacity. Therefore,two different models were developed to reflect the types of gas turbine systems employed inthe plant models: one for a large gas turbine system with a capacity of over 300 MWel forthe IGCC with entrained flow gasification, and one for a medium-size gas turbine systemwith a capacity of 30–70 MWel.

The gas turbine systems are characterized by their compressor pressure ratio, combustoroutlet temperature15, and isentropic efficiencies of compressor and expander. Compressionand expansion are modelled in four stages, whereat cooling air can be extracted or addedat each stage. The first three expander stages are assumed to be cooled and total coolingair is assumed to be 10% of inlet air, based on the PG9351 FA turbine model in theGateCycle Gas Turbine Standard Library. The modelling assumptions, including heatand pressure losses, are given in Table A.24.

The efficiencies for natural gas operation are 35.2% and 31.8% (HHV), respectively, or39.0% and 35.3% (LHV). In Figure 3.5, the turbine inlet temperatures and efficiencies of15The turbine inlet temperature equals the combustor outlet temperature in the simulations.

69


25%

30%

35%

40%

45%

50%

0 100 200 300 400

rating [MWel]

elec

trica

l effi

cien

cy (L

HV)

literaturemodel

1000

1100

1200

1300

1400

1500

1600

0 100 200 300 400

rating [MWel]

TIT

[°C]

literaturemodel

Figure 3.5: Efficiency and turbine inlet temperature (TIT) of gas turbine systems in re-lation to their capacity. Data from [301] and the gas turbine models used in this work.

Table 3.8: Steam cycle parameters.

plant type large IGCC medium IGCC CHPboiler type HRSG HRSG wood-firedsteam turbine capacity [MWel] >300 70–90 3–25cycle type 3 pressure reheat 2 pressure reheat 1 pressuresteam pressure levels [bar] 127/41/7 110/6 80live/reheat steam temperatures 1) [°C] ΔTmin=20 ΔTmin=20 500temperature difference at pinch point [°C] 10 10 n.a.condenser pressure [bar] 0.05 0.05 > 1.0

1) ΔTmin is the gas inlet/steam outlet temperature difference

the models used here and several existing gas turbine systems based on data from [301]are shown in relation to their capacity (rating).

The same gas turbine simulation models are then employed fuelled with hydrogen-richsyngas in the IGCC plant configurations described in chapter 5. The syngas is assumed tobe diluted with nitrogen and water vapour to a hydrogen content of 50% before combustion,for reasons of NOX control and flame stability [263].

The exhaust gas to inlet air ratio is higher for a gas turbine run on hydrogen-rich syngasrelative to natural gas operation. This might require changes to the geometry or operatingconditions, such as pressure ratio or guide vane angles, for existing turbine models [302].Such issues were not considered in this work.

3.3.2.15 Steam cycle

Steam cycles are employed as the bottoming cycle in the IGCC plants and in the wood firedCHP plants. The key design parameters are given in Table 3.8. Assumptions regardingsteam turbine efficiencies, pressure losses and so forth are given in sections A.2.1.2 toA.2.1.4.

70

3.4 Economic assessment

3.3.2.16 Air separation unit

The cold box of the cryogenic air separation unit (ASU) is modelled as a black box,characterized by its heat and electricity consumption and its operating pressure, based ondata from literature. The air, oxygen and nitrogen compressors are modelled as intercooledcompressors with 2–6 stages as described in section A.2.1.2.

For the IGCC with entrained-flow gasification, the ASU is operated at an elevated pressureand receives 29% of its air supply from the gas turbine compressor.

Details on the modelling assumptions for the ASU are given in section A.2.1.13.

3.3.2.17 Reciprocating engine CHP module

A biogas-fired reciprocating engine is employed in simulation case ADP-1. The engine ismodelled on the 600 kWel module TCG 2016 V12 C (Biogas) by German manufacturerMWM. Details are given in section A.2.1.14.

3.3.2.18 Refrigeration machine

In the anaerobic digestion plants, a refrigeration unit is used to condense water fromthe biogas, in order to fulfill the requirements of downstream plant equipment or meetthe quality standards for injection into the natural gas grid. Tetrafluoroethane (C2H2F4,R134a) is used as the refrigerant to cool the biogas to 3°C. The refrigeration unit modelcomprises a condenser, evaporator, compressor and throttle valve. The pressure is 10 barin the condenser and 2 bar in the evaporator.

3.3.2.19 Miscellaneous auxiliary units

Auxiliary units not included in the simulations (and not contained in the flowsheets)comprise fans for cooling air, cooling water pumps and miscellaneous other equipment(including water demineralization, waste water treatment, slag treatment, control systems,lighting). Their electricity consumption is accounted for in the energy balance of theplants, as detailed in section A.2.1.15.


Based on the simulation results, the investment, feedstock and auxiliary energy costs areestimated, and the annual levelized product costs are then calculated. The analysis isbased on constant year 2010 € (without inflation). The interest rate, capacity factor16

and economic plant lives are given in Table 3.9.

16The capacity factor is the actual annual output divided by the potential output when run at nameplatecapacity for 8760 hours.

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Table 3.9: Key assumptions for the economic assessment.

real interest rate [% p.a.] 10.0%capacity factor [–] 80%economic plant life

upgrading plant [a] 15CHP plant [a] 15large power plant [a] 30

3.4.1 Estimation of the capital investment

The level of confidence of an investment cost estimate depends on the level of detail inthe plant design and on the maturity of the technologies involved [303]. Study estimatesapplying scaling exponents to previously established plant cost usually exhibit an uncer-tainty range of ±20–30% [304, page 15]. However, the uncertainty increases dramaticallyfor technologies in development status, ranging from −30% to +50% for technologies inpilot stage and −30% to +200% for processes proven only at laboratory scale [303, pages1–6]. Most technologies considered in this work contain at least some components in pilotstage, thus the uncertainty of the investment cost estimates is substantial.

The cost estimates presented here do not contain any R&D expenditures or risk premiumsfor a first-of-a-kind plant. For the technologies still at a pilot stage, these estimates there-fore represent future commercial plants after some demonstration projects have provedsuccessful.

Cost data is converted to year 2010 € using inflation indexes for chemical engineeringplant equipment and adjusted to the respective sites with location factors for Germanyand Malaysia. For details, see section A.3.1.

3.4.1.1 Plant equipment

For standard chemical plant equipment, such as heat exchangers, pumps and processvessels, cost data is mostly taken from chemical engineering handbooks [304–308]. Formore specialized biomass processing equipment, such as HTC slurry pumps, pellet pressesand wood chips driers, cost estimates are based on vendor information. For the BECCSplants, some data had to be taken from cost estimates from other studies because no otherdata was available.

All plant equipment is sized based on its capacity Xsim resulting from the simulation. Asrecommended by Peters et al. [308, page 83], an overdesign factor fd of 1.1–1.2 is appliedto most equipment in the biomass upgrading plants. This should avoid bottlenecks in theprocess when operating conditions deviate from the design point, especially in regard tothe feedstock quality. The factor fd also accounts for spare parts, if required for safetyreasons or because of low availability of the equipment.

Thus the design capacity Xd is

Xd = fdXsim (3.9)

72


Xd may be split between n units to facilitate part load operation or because of a limitedmaximum capacity per unit.

Xunit =Xd

n(3.10)

The module cost per unit CBMunit is estimated with the help of dedicated cost functionsor more generally with a scaling exponent α:

CBMunit = CBMref ·(

Xunit

Xref

)α

(3.11)

where Xref is the capacity and CBMref is the cost of the reference unit for which the costis known.

The total module cost for a given process unit is

CBM = n · CBMunit (3.12)

Values for fd and n for the IGCC plants are based on [246].

The module cost CBM comprises the purchased equipment cost (PEC), equipment erec-tion, piping, instrumentation, electrical installation, process buildings, design and engin-eering.

If only the PEC is given in the data source, a factor for CBM/PEC is estimated forthe respective type of equipment. CBM/PEC is typically 2.5 for equipment processingsolids and 3.3 for equipment processing fluids [306, page 252].17 For power plant equip-ment, CBM/PEC is mostly in the range of 1.1–2.2,18 because instrumentation, electricalinstallation and so forth is included to a greater degree within the vendor’s scope.

At times, cost data is available only for lower specifications, i.e. for equipment made ofcarbon steel (CS) and operating below 10 bar. Material factors fM are applied to accountfor more expensive materials, typically stainless steel (SS). Pressure and temperaturefactors fp and fT account for more demanding operating conditions.

A comprehensive list of all the cost data and sources used for this work is presented inTables A.27 to A.44.

Especially for the most expensive items of each plant design, two or more different costfunctions are used when available. The arithmetic mean of the different cost functionsis then used as the CBM estimate. This approach is meant to lessen the risk of severelyunder- or overestimating importing equipment items. The results from the individualcost functions are presented19 for selected plant equipment to illustrate the degree ofuncertainty in the cost estimates.

17The factors given in [306] have been adjusted to the scope for CBM as mentioned above, excludingutilities, storages, site development, ancillary buildings, but including design and engineering.

18Based on the ratio of erect cost plus engineering plus accessories to the equipment cost of the mainequipment item, using data from [309].

19Refer to the tables headed “equipment list with investment costs” for the various plant designs given inthe appendix.

73


Table 3.10: Assumptions for the calculation of the total capital investment (TCI).

fees & contingencies 15% of CBMstart-up cost 1 month maintenance cost, 1 month labour, 1 week fuel, 2% of FCI

(without land)working capital 125% of the sum of 2 month maintenance cost, 2 months fuel, and 3

months labourdeconstruction cost 5% of FCIconstruction times

upgrading plant: 2 years large coal-fired power plant: 4 yearsCHP plant: 2 years large gas-fired power plant: 2 years

3.4.1.2 Total capital investment (TCI)

The fixed capital investment (FCI) includes the total module costs (CBM) plus fees andcontingencies and offsite costs.

The offsite costs comprise land, ancillary buildings, site development and utilities. For eachtechnology, these are estimated for one selected reference plant as a percentage of PEC. Forall other cases using the same technology, they are then calculated based on the capacityand a scaling exponent α according to Equation 3.11. For details see section A.3.3.

The total capital investment (TCI) comprises the FCI plus working capital, start-up costs,and allowances for funds used during construction (AFUDC), less the net present valueof the residual value (i.e. the revenue from reselling the land and recovery of the workingcapital minus the deconstruction cost).

Start-up cost and working capital are calculated with a rule of thumb estimate for electricpower plants, according to [310] (cited in [311]).

The AFUDC are calculated assuming that all payments are made as a lump sum at themidpoint of the construction period, as suggested by [312, page 17].

The assumptions for calculating the TCI are summarized in Table 3.10.

3.4.2 Carrying charges

The annual carrying charges are the obligations associated with the capital investment.They comprise the annuity of the TCI plus taxes and insurances to the amount of 1% ofthe TCI.20

3.4.3 Feedstock, auxiliary energy and other consumables

The feedstock cost at the upgrading plant gate comprises the cost at farm gate (seeTable 3.3) plus transport and storage as described in section 3.2.

Auxiliary energy for the upgrading plants comprises electricity, as well as oil palm shellsas a boiler fuel in the case of HTC from EFB. All other upgrading plants use part of the20The definition of carrying charges used in this work is based on [311, page 374 f.].

74


feedstock or biofuel as their boiler fuel, thus the boiler fuel is automatically included inthe feedstock cost.

Other consumables include chemicals and the disposal of waste streams such as ash, wastewater, and AD digestate.

The cost of electricity and other consumables are listed in Table A.48.

3.4.4 Labour

The labour requirement depends on the degree of automation and the susceptibility of theplant to blockage and similar. Since there is no experience with the continuous operationof most of the analyzed processes, it is difficult to assess the labout requirement with anydegree of certainty. The estimate for the upgrading plants is based on data for pelletizingplants [143, 313], wood-fired CHP plants [314] and a feasibility study for a fast pyrolysisplant with a capacity of 50 MW [315]. The labour requirement for the IGCC plants isbased on [316].

In analogy to the investment cost, the labour requirement is estimated for a referenceplant of each technology and adjusted to the plant capacity with a scaling exponent asdescribed by Equation 3.11. A minimum of one operator per shift is required at all times.

The annually required labour hours are estimated for plant operators and for workershandling the biomass in the preparation yard and the product. The annual labour costsare then calculated with the appropriate hourly rates. Administration, distribution andmarketing are not taken into account.

Details of labour costs are given in section A.3.5.

3.4.5 Maintenance material

Maintenance materials and spare parts for the upgrading plants are assumed to be 10%of the CBM for components with high wear and 2% of the CBM for all other components.The high wear components are mainly equipment items with moving parts processingabrasive solids. They include the HTC slurry pumps, screw presses, filter presses, pelletpresses, wood mills and the engine CHP module. For the IGCC plants, 2% of the CBMis assumed for all equipment items.

3.4.6 Carbon certificates

Making robust predictions on the development of the CO2 price over an economic plantlife of 15 or 30 years is next to impossible. The influence of CO2 price on the economicviability of the plants is therefore analyzed using sensitivity analysis. The product costsreported in this work do not contain CO2 costs unless otherwise stated.

The GHG mitigation cost defined by Equation 3.6 refers to the total amount of net(avoided) GHG emissions. These include cultivation and harvest, transport, and the

75


upstream emissions for auxiliary fuels. Due to the statuatory framework of emission trad-ing schemes, not all the GHG emissions are necessarily covered. The European UnionEmission Trading Scheme (EU ETS) in phase II only covers direct CO2 emissions.

Some of the observations on the economic viability of upgraded biofuels discussed insection 4.6.3 are related to the CO2 price for emissions covered by the ETS. They arereferred to as ETS carbon certificates to distinguish them from the total GHG emissions.

3.4.7 Levelized product cost

The annual levelized costs for auxiliary energy, raw materials, and operation and main-tenance are calculated with the constant escalation levelization factor (CELF) accordingto Equations 3.13 and 3.14.

CELF =k(1 − kn)

1 − kCRF with k =

1 + r

1 + i(3.13)

Cl = CELF · C0 (3.14)

where Cl is the annual levelized cost, C0 is the annual cost at the start of the project, CRFis the capital recovery factor, i is the annual interest rate and r is the annual escalationrate for the respective type of expense.

The escalation rates r used for fuels, electricity and other commodities are given inTable A.52.

The levelized product costs are the sum of the carrying charges and all levelized expenses,minus the revenues from byproducts.21

3.5 Exergy and exergoeconomic analysis

The exergy concept combines information from the first and second law of thermody-namics, and thereby allows one to calculate the thermodynamic value of material andenergy flows, and to assess the thermodynamic inefficiencies occuring within the plantcomponents.

Important parameters in exergy analysis are the rate of exergy destruction ED,k occurringin each plant component k and the exergetic efficiency εk. The exergy destruction iscalculated from the exergy balance

ED,k =∑

Ein,k −∑

Eout,k (3.15)

where Ein,k and Eout,k are the exergy flows related to material or energy streams enteringand exiting the component. The calculation of the exergetic efficiency requires a definitionfor the fuel EF,k and product EP,k of the component in terms of exergy. Exergy flows21The only byproduct considered in this work is electricity from biomass-fired CHP plants and some

upgrading plants.

76


exiting a component can also be classified as exergy losses EL,k, if they have no furtherutility. The definition of waste streams is usually restricted to the exergy balance of theoverall plant. On a component basis, all exergy flows are generally assigned to the fuel orproduct definitions, and EL,k is zero.

εk =EP,k

EF,k

=1 − ED,k − EL,k

EF,k

(3.16)

The relative impact of the plant components on the overall process can be expressed interms of the exergy destruction ratio, which relates the exergy destruction of a plantcomponent to either the fuel exergy EF,tot or the exergy destruction of the overall plant,ED,tot:

yD,k =ED,k

EF,tot

(3.17)

y∗D,k =

ED,k

ED,tot

(3.18)

Exergoeconomic analysis combines exergy analysis with cost calculations. By assigningcosts to the exergy flows within the plant, one can allocate monetary values to the ther-modynamic inefficiencies.

The cost of the exergy flows are calculated based on cost balances for the plant components:

∑Cout,k =

∑Cin,k + Zk (3.19)

where Cin,k and Cout,k are the cost associated with the exergy flows entering and leavingcomponent k, and Zk is the cost associated with the investment and maintenance of theequipment item.

The cost per unit of exergy ci for the flow stream i is defined as

ci =Ci

Ei(3.20)

Components with n outlet streams, in addition to the cost balance given in Equation 3.19,require n−1 auxiliary equations to allocate the costs to the various outlet streams. Theseauxiliary equations are related to the definition of the fuel and product streams of therespective components. Fuel and product definitions and auxiliary equations for commonplant equipment are given in [311, 317]. The rationale behind these definitions is explainedin [318].

The cost effectiveness of component k can be assessed with the help of the cost rate ofexergy destruction CD,k, the relative cost difference rk between the cost per unit of fueland product exergy cf,k and cp,k, and the exergoeconomic factor fk.

CD,k = cf,kED,k (3.21)

77


rk =cp,k − cf,k

cf,k(3.22)

fk =Zk

Zk + CD,k

(3.23)

These exergoeconomic variables provide insight into the cost formation process within theplant and can aid the discovery of potentials for cost reductions in the plant design.

Further background on the method of exergoeconomic analysis is presented in [311, 319].

3.5.1 Application of exergy-based analysis to biomass upgrading

Exergy analysis is applied to all the plant configurations analyzed in this work, whileexergoeconomic analysis is applied to the HTC base design only.

In order to allow a more resolved definition of fuel and product, exergy is commonlysplit into physical (PH ) and chemical (CH ) exergy. It has also been proposed to furthersplit physical exergy into mechanical and thermal contributions and chemical exergy intoreactive and non-reactive contributions [311].

In the analysis of the upgrading plants, the chemical exergy is the sum of the chemicalexergy of the solid biomass or biofuel dry matter, subscripted s, and the chemical exergyof the liquid and gas phase, subscripted lg.22

Etot = mseCHs + mlgeCH

lg + mtoteP Htot (3.24)

Ctot = CCHs + CCH

lg + CP H (3.25)

This facilitates a definition for the exergetic efficiency of the HTC reactor where the actualupgrading, i.e. the increase in quality of the feedstock, is defined as the product exergy:

εreactor =ms,out

(eCH

s,out − eCHs,in

)+ EP H

out + EP Hgas − EP H

in

(ms,in − ms,out) eCHs,in + ECH

lg,in − ECHlg,out − ECH

lg,gas

(3.26)

where the indices in and out indicate the biomass and biocoal slurries, respectively, andgas denotes the gaseous byproducts. This efficiency definition acknowledges that only partof the feedstock, ms,out, is upgraded, while the rest is consumed in chemical reactions.

Several options exist to define the exergetic efficiency of the overall upgrading plants. Thesimplest is to define the biofuel as the product and the raw biomass feedstock, electricity,auxiliary fuel (denoted by the subscript auxfuel), air for combustion and drying and make-

22Note that the biomass or biofuel moisture is included in the liquid and gas phase, not in the solid phase.

78


up water as the fuel.

εItot =

Ebiofuel

Ebiomass + Wel + Eauxfuel + Eair + Ewater(3.27)

The advantage of this definition is that it can be applied to all the analyzed technologies,thereby allowing comparisons to be made.

For HTC and torrefaction, the exergetic efficiency can alternatively be defined by analogyto that of the HTC reactor (Equation 3.26):

εIItot =

ms,biofuel

(eCH

s,biofuel − eCHs,biomass

)+ EP H

biofuel

(ms,biomass − ms,biofuel) eCHs,biomass + EP H

biomass + Wel + Eauxfuel + Eair + Ewater

(3.28)

In the scenarios considered in this work, EP Hbiomass and EP H

biofuel are zero, because the biomassis delivered at ambient conditions, and the biofuel is cooled down to ambient temperatureduring storage. For wood pelletizing, the efficiency εII

tot defined by Equation 3.28 cannotbe defined, because there is no increase in the specific chemical exergy of the wood drymatter.

For anaerobic digestion, it can be argued that the process “skims” part of the organicmatter from the feedstock to convert it to biomethane, while the digestate remains foragricultural applications. Following this rationale, the exergy of the converted materialshould be regarded as part of the fuel for the process:

εIIItot =

Ebiofuel(Ebiomass − Edigestate

)+ Wel + Eauxfuel + Eair + Ewater

(3.29)

For the analyzed technologies, results are given for each applicable efficiency definition.Definitions for exergetic efficiencies and auxiliary costing equations for equipment notdiscussed in [311, 317] but developed as part of this project are given in section A.4.

For the HTC reactor processing waste biomass, the cost rate of exergy destruction accord-ing to Equation 3.21 results in a negative value due to the negative cost of the biomass(meaning that the biomass supplier pays for the treatment). This would suggest that themore exergy destroyed, the better. However, this is not true, since more exergy destructionmeans that less biocoal can be generated from a given amount of biomass. Therefore, thecost rate of exergy destruction for the HTC reactor processing waste biomass at negativecost is calculated based on the cost per unit of product exergy cp,k rather than fuel exergycf,k, namely:

CD,k = cp,kED,k (3.30)

79


3.5.2 Application of exergy analysis to the BECCS plants

A BECCS plant has two purposes: the provision of energy in a desired form, eitherelectricity or syngas for the processes considered in this work, and the removal of CO2from the atmosphere. Therefore, the CO2 stream compressed for transport and storage,ECO2, is considered as part of the exergetic product. The exergetic efficiency of an IGCCplant is therefore defined as

ε =Wel,net + ECO2

Efeed + Eair + Ewater + Esand + Edolomite

(3.31)

where Ewater is the feedwater make-up compensating for water losses and Esand andEdolomite cover bed material losses from the fluidized bed gasifier.

3.6 Definitions for yields and efficiencies

Besides the various exergetic efficiencies, some energy related yields and efficiencies arealso used. Their definitions are given in the following.

Torrefaction and HTC are often evaluated using energy, mass and carbon yields. Theenergy yield γe is defined as the fraction of the chemically bound energy of the feedstockwhich is retained in the upgraded biofuel, without taking into account the auxiliary energyconsumption of the process.

γe =mbiofuelHHVbiofuel

mbiomassHHVbiomass(3.32)

Similarly, the mass yield γm and the carbon yield γc are the fractions of the dry feedstockmass and the share of the feedstock carbon retained in the upgraded biofuel, respectively.The carbon yield is especially relevant for BECCS applications, because only the carbonretained in the upgraded biofuel can later be captured at the BECCS plant.

The equivalent to the energy yield for gasification is the cold gas efficiency:

CGE =msyngasHHVsyngas

mfeedHHVfeed(3.33)

In contrast to the energy yield γe, the energetic efficiency ηHHV takes into account the aux-iliary energy consumption of the upgrading process in the form of electricity and auxiliaryboiler fuel.

ηHHV =mbiofuelHHVbiofuel

mbiomassHHVbiomass + Wel + mauxfuelHHVauxfuel

(3.34)

Some upgrading processes produce electricity and/or thermal energy Q as byproducts, inwhich case the efficiency becomes

ηHHV =mbiofuelHHVbiofuel + Wel + Q

mbiomassHHVbiomass + mauxfuelHHVauxfuel(3.35)

80

3.6 Definitions for yields and efficiencies

where Wel is the net electricity production (after subtracting the demand of the upgradingplant).

The higher heating value (HHV) is used throughout this work unless otherwise stated.

The lower heating value (LHV) does not account for all the energy in the fuel but onlythe share that can be recovered in a combustion facility without flue gas condensation.Somewhat paradoxically, LHV energetic efficiencies ηLHV can be greater than 100% forcombustion facilities with flue gas condensation, and for upgrading processes like HTCwhere moisture is removed from the fuel in liquid form. An ηLHV greater than 100% foran upgrading process indicates that more energy can be potentially recovered from burningthe upgraded biofuel than from burning the raw biomass, when using a combustion facilitiywithout flue gas condensation.

The upgrading processes consume electricity and biomass in various ratios. Comparedto biomass, electricity is a more refined form of energy which has been produced in alossy conversion process. Therefore, it is common to take into account the primary energyfP E

˙·W el required to provide the electricity. This leads to a modified efficiency definition:

ηLHV,P E =mbiofuelLHVbiofuel

mbiomassLHVbiomass + fP E˙·W el + mauxfuelLHVauxfuel

(3.36)

A primary energy factor fPE of 2.67 GJPE/GJel is used in this work based on the EU-17power mix [320].

The conversion chain efficiency from raw biomass to electricity is defined as

ηCC,HHV =Wel,P P − Wel,UP G

mbiomassHHVbiomass + mauxfuelHHVauxfuel(3.37)

where Wel,P P is the net electricity production of the power plant and Wel,UP G is theelectricity consumption of the upgrading plant. The captured CO2 cannot be accountedfor as a product of a BECCS process in an energy-based efficiency definition.

For the IGCC plants analyzed in section 5.2, gross and net efficiencies can be defined,depending on whether the electricity consumption of the process units is accounted for:

ηgross,HHV =Wel,produced

mfuelHHVfuel(3.38)

ηnet,HHV =Wel,produced − Wel,consumed

mfuelHHVfuel(3.39)

The carbon capture rate for the conversion chain is defined as

γC,CC =mC,captured

mbiomasscbiomass + mauxfuelcauxfuel(3.40)

where mC,captured is the carbon in the captured CO2 stream, and the c are the carbonmass fractions of the feedstock and the auxiliary biomass fuel. A small fraction of the

81


feedstock carbon ends up in ash or slag and may be stable on a long term, depending onthe form of disposal. However, this is not accounted for as captured carbon.

82

4 Biomass upgrading processes

This chapter considers a number of processes for upgrading biomass. Since drying con-tributes significantly to the energy consumption of all the considered processes generatingsolid biofuels, different drying technologies are first analyzed and compared. Subsequently,plant designs for wood pelletizing, torrefaction, anaerobic digestion and HTC are presen-ted. Their performance is evaluated in terms of energetic and exergetic efficiency, GHGemissions and cost. The different biofuels are then discussed and compared as substitutesfor fossil fuels in existing power stations.

4.1 Drying

The energy consumption for drying has a major impact on the efficiency of biomass conver-sion technologies. The share of the feedstock energy (HHV) required for the evaporationof the moisture is 12.5% for wood with a water content of 50%, without accounting forthe energy losses of the drier. For waste biomass with a water content of 70%, 33–40% ofthe feedstock energy is consumed by the evaporation of the moisture. This indicates thatfor wet feedstocks, conversion technologies which do not require prior drying, as well asdrying technologies which allow the enthalpy of evaporation to be recovered, are especiallyattractive. In the case of HTC, only 6.6% of the biocoal energy is required to evaporatethe remaining moisture following mechanical dewatering of the biocoal to a water contentof 40%.All technologies for the production of a solid biofuel require drying either before theconversion process or afterwards. Because of the large impact of drying on efficiency,simulations in Aspen Plus and exergy analysis for three different drying technologies areperformed for wood and biocoal.The following drying systems are analyzed:Low temperature drying (LTD) uses hot water, steam or other relatively low temper-ature heat sources to heat up the drying air. Belt driers are commonly used for drying attemperatures between 75–110°C [28]. This drying system is attractive if low temperaturewaste heat is available, for example the exhaust steam from the steam turbine in a rankinecycle or the waste steam from biocoal slurry de-pressurization in an HTC plant.Simulation case W-LT-1 represents the drying of wood chips with steam at 1 bar, W-LT-2represents the drying of wood chips with steam at 10 bar, and BC-LT-1 models the dryingof biocoal with steam at 1 bar.High temperature drying (HTD) uses combustion gases as the heat source. Rotarydrum driers operated on hot combustion gases with temperatures up to 600°C are increas-ingly employed in wood pellet factories due to their compact design and high performance[28].

83

Chapter 4 Biomass upgrading processes

High temperature drying of wood chips is modelled in simulation case W-HT. Raw woodis burned to produce the hot gas required for the drier.1 Part of the cooled combustiongas is recirculated to limit the temperature of the gas at the drier inlet to 500°C.

Superheated steam drying (SSD) allows the recovery of the heat of evaporation. Sincethe drying in steam atmosphere prevents the dilution of the evaporated moisture with air,its enthalpy of condensation is available in the form of steam at 100°C if the drier isoperated at atmospheric pressure. Operating the drier at an elevated pressure allows forthe recovery of steam at higher temperatures, thus facilitating its integration in adjacentprocesses. If there is no demand for the recovered steam in adjacent processes, it can bere-compressed to provide superheated steam for the drier.

cond.

cond.

wetbiomass

wetbiomass

wetbiomass

wetbiomass

wetwood

exhaustgas

exhaustgas

driedbiomass

driedbiomass

driedbiomass

driedbiomass

SSD

SSD

air

air

a) c)

d)b) steam

steam

steam

condensate1

1

91

1

6

6

8

9

7

710

3

3

3

8

8

10

4

4

6

6

7

7

5

5

11

11

12

1313

8

15

2

2

2

2

4

4

1411 12

drier

drier ash

W1

W1

W1

W2

W2

W2

W1

Figure 4.1: Drying systems: a) high temperature drying (W-HT ), b) low temperaturedrying (W-LT-1, W-LT-2, BC-LT-1 ), c) SSD for wood chips (W-SSD-1 ), d) SSD forbiocoal (BC-SSD-1, BC-SSD-2 ).

High temperature rotary drum driers and low temperature belt driers are both commonlyused, while SSD remains limited to about 100 large-scale applications [321]. Both atmo-spheric and pressurized fluidized-bed drying in a steam atmosphere have been applied tolignite in pilot and demonstration plants [322, 323]. Atmospheric fluidized-bed driers forlignite are usually operated with a small particle size of 0.2–6 mm, because a finer particlesize significantly improves the heat transfer and fluidization properties [186, 291]. Thehigher pressure in pressurized SSD driers leads to higher heat transfer coefficients andallows a more compact design [322]. Pressurized SSD has been applied to wood chips andsaw dust in Sweden [293, 324]. The operation was not without problems, due to issues withthe feeding systems and the disposal problems of terpene containing condensate [293, 325].

Simulation case W-SSD-1 represents the pressurized SSD of wood chips. The drier oper-ating pressure is 4 bar. Part of the drier exhaust steam is recompressed to 4.5 bar andsupplied as fluidization steam to the drier. The remainder is compressed to 15 bar and

1Assumptions for the combustion: stoker boiler with 3% unburned carbon, 3% radiative heat loss, airratio 1.4

84

4.1 Drying

used as heating steam for the drier. The condensate is then flashed to 4 bar to recover moresteam. Some additional steam at 15 bar (11)2 is required to fulfil the energy demand of thedrier. Case BC-SSD-1 models a similar design for biocoal, but in contrast to W-SSD-1,the condensate (8) is utilized here to preheat the wet biocoal. Case BC-SSD-2 representsthe same design as BC-SSD-1, but the drier is operated at atmospheric pressure, and theheating steam is provided at 3.4 bar.The flowsheets of the respective systems are shown in Figure 4.1. All simulation casesmodel drying to a final water content of 10%. The results of the exergy analysis aresummarized in Table 4.1. Only steam, combustibles and electricity are classified as fuelin Table 4.1, since these are the exergy flows relevant for a techno-economic comparison.There is also a considerable chemical exergy input from the inlet air, particularly with LTdrying.

Table 4.1: Exergetic fuel and exergy destruction and exergy losses of different dryingsystems for wood chips and biocoal. All values are normalized to 1 kg of evaporated(removed) water in [kJex/kgH2O,ev].

W-LT-1 W-LT-2 W-HT W-SSD-1 BC-LT-1 BC-SSD-2 BC-SSD-1

fuel

steam 740 956 — 222 752 82 305

combustibles — — 3711 — — — —electricity 197 109 27 529 202 488 368

total 937 1065 3738 751 954 570 674

inlet air and moisture

inlet air 210 115 7 — 214 — —inlet moisture 56 56 56 57 60 60 60

exergy destruction

drier 501 528 1267 485 510 355 362

boiler — — 1798 — — — —heat exchanger 345 391 — — 350 61 92

fan, compressors 56 32 2 112 59 118 88

total 901 951 3067 598 919 533 542

exergy losses

waste heat 1) 55 124 561 116 54 20 97

ash — — 106 — — — —other 248 161 66 95 255 77 95

1) EPH of exhaust gas and condensate

In case W-HT, with high temperature drying, the total exergy demand of steam, combust-ibles and electricity is four times as high as that of case W-LT-1. The exergy destructionin the HT drier itself is 2.5 times as high as that within the LT drier, and the additionalexergy destruction in the boiler is even higher than that of the drier. This clearly showsthat using waste heat instead of burning fuel to provide thermal energy to the drier shouldbe pursued whenever possible. The physical exergy of the exhaust gas is higher than inthe cases with LT drying. This exergy will be lost to the environment in most cases butfurther utilization is conceivable where there is a demand for low grade heat.

2Numbers in parentheses refer to the flowstreams on the relevant flowsheet.

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LT drying requires much bigger air flows than HT drying to remove the feedstock moisture,since the saturation pressure of cold air is lower. This reflects in the higher electricityconsumption of the fan.

In case W-LT-2, the fuel exergy is 14% higher than in W-LT-1 due to the higher specificexergy of the steam at 10 bar. The electricity consumption of the fan is lower in W-LT-2because less drying air is needed at a drying temperature of 142°C, compared with 90°Cin W-LT-1.

Pressurized SSD reduces the required fuel exergy by 20–30% compared to LT drying,and atmospheric SSD reduces it by 40%. Though less efficient than atmospheric SSD asa stand-alone application, pressurized SSD offers the advantage of making the exhauststeam available at a higher temperature and pressure. Depending on the possibilities forprocess integration, this may eliminate the need to re-compress the exhaust steam. Thefuel exergy of W-SSD-1 is only 20% that of W-HT.

The fuel exergy per kg of evaporated water is comparable for wood chips and biocoal.However, the amount of water that needs to be evaporated per kg of dry matter is 37%lower for the biocoal than for wood chips. This, together with the higher calorific valueof the biocoal, results in the fuel exergy per MJHHV of dried biomass being 50% lower forbiocoal than for wood.

Due to its superior performance, SSD is integrated into several plant designs analyzed inthe following sections. The base designs, however, employ the more proven HTD and,where waste heat is available, LTD technologies.

4.2 Pelletization

Wood pelletization serves as a reference process for the other upgraded biofuels tech-nologies processing wood. Energy balance, carbon balance and economic analysis areperformed for a simple pelletizing plant, a plant design with SSD and a plant design withan integrated CHP process.

4.2.1 Design and simulation models of a wood pelletizing plant

Simulation case WP-1.0 represents a simple wood pelletizing plant with high temperaturedrying. Thermal energy is provided to the drier by the combustion gases derived fromburning raw wood. The simulation model is rather simple, comprising the unit operationsdrying, wood combustion in a stoker boiler to provide thermal energy for the drier, millingand pelletizing. Simulation case WP-1.1 models a wood pelletizing plant with pressurizedSSD, where the drier exhaust steam is re-compresssed to provide fluidization and heatingsteam to the drier. Additional steam required to fulfill the energy demand of the drier isprovided by a boiler fuelled on raw wood.

Case WP-1.2 models a wood pellet factory with a CHP plant consisting of a wood boilerand simple back-pressure steam turbine. The flowsheet is shown in Figure 4.2. The steamturbine inlet parameters are 500°C and 80 bar (17). The exhaust steam from the steam

86

4.2 Pelletization

turbine at 1 bar (18) is utilized to supply thermal energy to the low temperature woodchip drier. The boiler exhaust gas (14) leaves the plant at 120°C.

Case WP-1.3 uses the same design as WP-1.2 but with a boiler exhaust gas temperatureof 50°C resulting in partial exhaust gas condensation.3

Cases WP-1.0 and WP-1.1 are designed for 2.0 t/h dry matter of raw material for thepellet production. In cases WP-1.2 and WP-1.3, the capacity is 10.0 t/h dry matter.

wetwood

wetwood

steamturbine

exhaustgas

exhaustgas

driedwood

air

airash

5

13

31

14 12

16

1518

2

17

9

6

10

drier

W1W4 W5

W2

W3

pelletpressmill

G

Figure 4.2: Flowsheet of the wood pellet plant with CHP, WP-1.2.

4.2.2 Energy and carbon balance

The energy demand and energetic efficiency of the analyzed plant configurations are shownin Table 4.2.

Of the overall biomass supplied to the plants, 16–26% is combusted to provide the thermalenergy for the drying process, or for steam generation for the CHP plant in the cases WP-1.0, WP-1.2 and WP-1.3. In WP-1.2 with SSD, only 4% of the feedstock biomass isburned. However, the electricity consumption in WP-1.1 is more than twice that of WP-1.0 because of the steam compression.

In WP-1.1, the carbon yield is the highest at 96.5%, because the SSD with recompressionof the exhaust steam greatly reduces the energy demand of the drying process and relatedwood combustion. The energetic efficiency of this configuration is also the highest. At92.6%, the HHV-based efficiency ηHHV is 10% higher than that of case WP-1.0 withconventional low temperature drying. The LHV-based efficiency ηLHV is greater than100% because the enthalpy of condensation of the fuel moisture is utilized and the moistureis removed in liquid form.

The energetic efficiencies of the configurations WP-1.2 and WP-1.3 with CHP are lowerthan that of mere wood pellet production. However, these processes constitute a higherdegree of conversion. Some of the feedstock biomass is converted from wood to electricity,while in cases WP-1.0 and WP-1.1, raw wood is converted to wood pellets, which is arather shallow conversion. It is therefore more useful to compare the overall conversionchain efficiencies when the wood pellets are co-combusted in a coal-fired power station.

3The subcooled condensate (15) is delivered at 35°C.

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Table 4.2: Energy demand and energetic efficiency for the four wood pelletizing cases.

WP-1.0 WP-1.1 WP-1.2 WP-1.3capacitypellets mass flow (FM) [t/h] 2.222 2.222 11.111 11.111wood mass flow, total (FM) [t/h] 4.773 4.146 27.354 25.868biomass consumptionwood for pellets [MWHHV] 10.860 10.860 54.299 54.299wood for boiler [MWHHV] 2.100 0.395 19.436 15.932electricity consumption1)

hammer mill [MWel] 0.084 0.084 0.419 0.419pellet press [MWel] 0.111 0.111 0.556 0.556drier fan [MWel] 0.038 0.000 0.597 0.597steam compressors [MWel] — 0.278 — —feedwater pump [MWel] — — 0.062 0.062steam turbine [MWel] — — -3.055 -3.055total [MWel] 0.233 0.473 -1.422 -1.422efficiencieselectrical efficiency [–] — — 1.9% 2.0%energetic efficiency, HHV [–] 82.3% 92.6% 75.0% 78.9%energetic efficiency, LHV [–] 93.3% 104.4% 85.6% 90.0%energetic efficiency, LHV, PE [–] 90.0% 96.4%carbon yield [–] 83.8% 96.5% 73.1% 76.9%

1) Negative values indicate production

The results are shown in Table 4.3.4 WP-1.1 with the SSD has the highest overall con-version chain efficiency at 30.8%. Case WP-1.0 with conventional drying and WP-1.2with conventional drying and CHP are almost identical at 28.7% and 28.5%, respectively,while the efficiency in WP-1.3 is 1.4 percentage points higher. 55% of the water vapour inthe boiler exhaust gas is condensed in WP-1.3, leading to its improved performance whencompared to WP-1.2.

The plants with CHP produce 1.4 MW surplus electricity which can be fed into theelectricity grid. The consumption by hammer mill, pellet press, fans and pumps amountsto 56% of the gross electricity production. The electrical efficiency5 of the CHP plant is15.6% in WP-1.2 and 19.1% in WP-1.3. Reasons for the rather low efficiency include thelow steam parameters, simple design of the steam cycle and high outlet pressure of thesteam turbine. Only 2% of the feedstock energy supplied to the combined CHP and woodpellet plants is converted to net electricity, while 73–77% are converted to wood pellets.

4.2.3 GHG emissions

The supply chain GHG emissions of the wood pellets range from −3.8 to 12.6 kg/GJ.Emissions arising from the road transport of the biomass to the pellet plant increase with

4For a discussion of the power plant efficiency for wood pellet firing, see sections 3.2.5 and 4.6.2.5Defined as ηCHP = WST −Wpump

mboiler fuel·HHVwood.

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4.2 Pelletization

Table 4.3: Conversion chain efficiency from wood to electricity when the wood pellets areco-combusted in a coal-fired power plant.

WP-1.0 WP-1.1 WP-1.2 WP-1.3co-combustion in power plantpower plant efficiency (HHV) [–] 36.3% 36.3% 36.3% 36.3%wood pellets [MWHHV] 10.860 10.860 54.299 54.299electricity production [MWel] 3.942 3.942 19.710 19.710overall conversion chainraw wood [MWHHV] 12.960 11.255 74.264 70.623el. consumption WP plant [MWel] 0.233 0.473 -1.422 -1.422net electricity production [MWel] 3.709 3.469 21.132 21.132energetic efficiency (HHV) [–] 28.7% 30.8% 28.5% 29.9%

plant capacity and contribute 0.4–2.0 kg/GJ. Road transport of the pellets from the plantto the harbour and shipping contributes 3.9 and 3.5 kg/GJ, respectively, for the caseswhere the biocoal is shipped from overseas. The electricity consumption is responsible for3.8 kg/GJ in the base case WP-1.0 and 7.8 kg/GJ in WP-1.1. In the cases with CHP, thevalue is negative at −4.7kg/GJ, because of the credit for the surplus electricity fed intothe grid. This results in negative net emissions for wood pellets produced in Europe fromthe processes that embed CHP. A breakdown of the GHG emissions from the wood pelletproduction cases can be found in Table B.41.

4.2.4 Economic performance

The wood pellet production costs are calculated for the four plant designs. In the case ofthe plant configurations with CHP, only the medium and large capacities are considered.Cost data from literature for most of the equipment was only available for small andmedium pellet factories. There is therefore a higher level of uncertainty regarding theinvestment cost estimates for the large scale plants.

The investment and wood pellet production costs are summarized in Table 4.4. Detailson the equipment cost estimates can be found in section B.1.1.

Specific investment costs reported in literature vary widely from 66 to 372 €/kWbiofuel forplant capacities between 21–70 kt/a for simple plant designs [106, 112, 326]. The calculatedinvestment cost for case WP-1.0-s and WP-1.0-m, which correspond to this plant typeand capacity range, here result in 173–283 €/kW. The investment required for the largestpellet factory worldwide (as of 2012) with a production capacity of 750 kt/a is reportedto be 120 million €, corresponding to 206 €/kW [110].

The TCI for the most efficient plant design WP-1.1 is more than 60% higher than forWP-1.0. The resulting pellet production costs are 4–9% higher. The pellet productioncost of the plant designs with CHP are 3–11% higher than that for WP-1.0, whereas theinvestment is about double. The most dominant cost contribution is the SR wood used asthe feedstock, contributing 48–85% of the total cost of pellet production. Other significantcost components are carrying charges (8–23%), labour (3–15%) and electricity (6–13%).The overall production costs range from 8.5–11.4 €/GJ, or 150–200 €/t.

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The scenario in which forest residues in North America are used as feedstock, road trans-port to the harbour amounts to 0.97 €/GJ and shipping to 1.08 €/GJ. The results areshown in Table 4.5. The overall production costs comprise 6.5–10.2 €/GJ, or 115–180€/t. Sikkema et al. report pellet prices of 115–140 €/t between 2007–2010 for overseaspellets landed at a port in the Netherlands [10]. This is in good agreement with the costscalculated here for medium and large scale plant, which also lie between 115–140 €/t.

The relative merits of the different designs depend on the wood and electricity prices.SSD could become attractive at high wood prices and low electricity prices, while CHP isfavourable with low wood prices and high electricity prices.

Table 4.4: Investment and levelized wood pellet production costs using short rotationwood.

WP- 1.0-s 1.0-m 1.0-l 1.1-s 1.1-m 1.2-m 1.2-l 1.3-m 1.3-l

pellet production [kt/a] 15.6 77.9 389.3 15.6 77.9 77.9 389.3 77.9 389.3

wood consumption [kt/a] 33.5 167.3 836.3 29.1 145.3 191.7 958.5 182.3 911.5

TCI [M€] 3.08 9.43 31.39 5.19 15.43 19.79 69.70 19.52 68.64

specific TCI [€/kWbiofuel] 283.2 173.5 115.5 477.2 283.8 364.0 256.4 359.1 252.5

carrying charges [€/GJ] 1.59 0.97 0.65 2.67 1.59 2.04 1.44 2.01 1.41

labour [€/GJ] 1.61 0.61 0.24 1.53 0.58 0.80 0.30 0.79 0.30

electricity [€/GJ] 0.61 0.61 0.61 1.25 1.25 -0.60 -0.60 -0.60 -0.60

O&M, material [€/GJ] 0.30 0.20 0.13 0.42 0.25 0.28 0.21 0.28 0.21

other operating cost [€/GJ] 0.03 0.03 0.03 0.03 0.03 0.03 0.03 0.03 0.03

feedstock [€/GJ] 4.77 4.77 4.77 4.14 4.14 5.46 5.46 5.20 5.20

feedstock transport [€/GJ] 0.28 0.47 0.74 0.23 0.40 0.56 0.88 0.53 0.82

feedstock storage [€/GJ] 1.31 1.31 1.31 1.14 1.14 1.51 1.51 1.43 1.43

total pellets cost [€/GJ] 10.50 8.98 8.48 11.41 9.36 10.07 9.22 9.65 8.80

total pellets cost [€/t] 185 158 149 201 165 177 162 170 155

Table 4.5: Levelized wood pellet production costs using forest residues in North Americaand then shipping the pellets to Europe.

WP- 1.0-s 1.0-m 1.0-l 1.1-s 1.1-m 1.2-m 1.2-l 1.3-m 1.3-l

pelletizing [€/GJ] 4.05 2.33 1.56 5.81 3.60 2.44 1.27 2.40 1.24

feedstock [€/GJ] 2.06 2.06 2.06 1.79 1.79 2.36 2.36 2.24 2.24

feedstock transport [€/GJ] 0.27 0.46 0.71 0.23 0.38 0.55 0.85 0.52 0.80

feedstock storage [€/GJ] 0.19 0.19 0.19 0.17 0.17 0.22 0.22 0.21 0.21

pellets transport [€/GJ] 2.04 2.04 2.04 2.04 2.04 2.05 2.05 2.05 2.05

total pellets cost [€/GJ] 8.61 7.08 6.57 10.03 7.98 7.62 6.74 7.41 6.53

total pellets cost [€/t] 153 125 116 177 140 134 119 130 115

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4.3 Torrefaction

4.3 Torrefaction

The energy balance, carbon balance and economic performance are assessed for a torre-faction plant which produces torrefied pellets from raw wood.

4.3.1 Design and simulation model of a torrefaction plant

The selected flowsheet design is one suggested by ECN and is described in several publica-tions [39, 114, 117, 121, 122]. It employs a directly heated torrefaction reactor, where thethermal energy is provided to the reactor by recirculated and reheated torrefaction gas.

4.3.1.1 Simulation model of the torrefaction reaction

The torrefaction reaction is modelled as a black box. Yields of products and byproductsare based on measured data from laboratory-scale experiments for the torrefaction ofwillow wood at 250°C with a residence time of 30 minutes [141].

Byproducts in the simulation model comprise CO2, CO, acetic acid and methanol. Sincemethanol is not reported separately in [141], it is assumed that the mass ratio of methanolto acetic acid is 0.25, based on data from [114]. The reported measured data [114, 141]includes a certain amount of unspecified “other organics”, therefore the mass yields ofacetic acid and methanol in the simulation are adjusted to reach the reported mass yieldof torrefied wood. The composition of the wood from [141] is slightly different than thatused in this work. Nevertheless it is assumed that the same mass yields (daf) of byproductsare generated. It is also assumed that all S, N and ash remain in the torrefied wood. Themass yield and composition of the torrefied wood are then calculated by difference.

The mass yields of torrefied wood and byproducts, higher heating values of raw andtorrefied wood and the energy yield for the torrefaction simulation model are given inTable 4.6. The measured data on which the model is based is included for comparison.The HHV of both the raw wood and the torrefied wood is slightly higher in the simulationmodel due to the different composition of the wood used as the feedstock, but the energyyield is almost identical. The composition of the torrefied wood used in the simulationand from [141] is given in Table B.4.

4.3.1.2 Plant model

An Aspen Plus simulation model originally developed by Brachnarova [143] was refinedand adjusted to the general assumptions made in this work. The flowsheet is shown inFigure 4.3. Upstream of the torrefaction reactor, the fresh wood needs to be dried to awater content of 20% [122]. The torrefied wood is cooled to a temperature of 40°C andthen pelletized. The water content of the torrefied wood is assumed to be 3% [39]. Itis also assumed here that the torrefied wood is pelletized without adding moisture. Inpractice, the pelletization process may require moisture to reduce friction in the pellet die,resulting in a final water content of the pellets of about 7% [122]. The byproduct of thetorrefaction reaction, the torrefaction gas (14), is compressed to overcome pressure losses

91


Table 4.6: Mass yields, HHV and energy yield for the torrefaction simulation model (perkg dry feedstock), and the measured data from [141] and [114], which the model is basedon.

simulation [141] [114]feedstock wood willow wood cuttingsT [°C] 250 250 280t [min] 30 30 17.5torrefied wood [kg/kg] 87.26% 87.20% 87.50%water [kg/kg] 5.67% 5.70% 8.00%CO2 [kg/kg] 2.89% 2.90% 1.22%CO [kg/kg] 0.30% 0.30% 0.17%CH4, H2 [kg/kg] 0.01%acetic acid [kg/kg] 3.10% 2.10% 1.16%methanol [kg/kg] 0.78% 0.29%other organics [kg/kg] 1.80% 1.66%HHV, wood (d.b.) [MJ/kg] 19.548 18.970HHV, torr. wood (d.b.) [MJ/kg] 21.258 20.594energy yield, HHV [–] 94.9% 94.7% 94.9%

in the system. Part of it (15) is heated to a temperature of 540°C and recirculated to thereactor, where it provides the thermal energy required to reach the reaction temperatureof 250°C. The remainder (16) is combusted. The combustion gases (19) provide thermalenergy to heat the recirculated torrefaction gases and for the drum drier. Part of thedrier exhaust gas (7) is re-compressed and mixed with the hot combustion gas to limitthe temperature to 600°C to reduce material stress. Wood is used as an auxiliary fuel,because the torrefaction gas alone does not provide sufficient thermal energy to supplythe torrefaction reactor and the drier.

It is assumed that the wood chips can be torrefied without further comminution, and thatthe torrefied wood can be crushed in the pellet press without a previous milling step.

Flowstream data from the simulation is given in section B.2.1.

wetwood

torrefiedpellets

wetwood

exhaustgas

pelletpress

air

1

298

13 18

25

21

3

5

6

27

14

17

1516

24

26

20 19

7

drier reactor

ashW2

W1

W3

W4

Figure 4.3: Flowsheet of the torrefaction plant TOR-1.0.

92

4.3 Torrefaction

4.3.2 Energy balance, carbon balance and GHG emissions

The energy balance and efficiencies of the torrefaction plant are summarized in Table 4.7.The energetic efficiency is 79.4% (HHV), or 91.8% (LHV), or 89.8% (LHV, PE). This isin good agreement with the published results from simulation studies of a similar plantdesign, which report an efficiency of 91% (LHV, including utility consumption) [144], and92% (LHV, PE) [327].

15% of the wood supplied to the plant is combusted in the furnace, while 4% of the overallfeedstock energy is converted to thermal energy and byproducts in the torrefaction reac-tion. Both the energetic efficiency and the carbon yield are lower than for simple woodpelletizing with conventional drying (WP-1.0 ) because of the conversion losses accompa-nying the torrefaction reaction.

Supply chain GHG emissions of the torrefied pellets from SR wood range from 3.0 to4.2 kg/GJ, comprising 0.5–3.0 kg/GJ from road transport of the biomass and 2.5 kg/GJfrom electricity consumption. Use of overseas forest residues and shipping of the pelletsto Europe results in 10–12 kg/GJ, similar to the equivalent wood pellet scenarios. Abreakdown of the GHG emissions for the torrefaction cases can be found in Table B.41.

Table 4.7: Energy demand and energetic efficiency of the torrefaction plant.

TOR-1.0capacitypellets mass flow (ar) [t/h] 1.799wood mass flow, total (ar) [t/h] 4.728biomass consumptionwood for pellets [MW] 10.86wood for boiler [MW] 1.98electricity consumptionpellet press [MW] 0.090compressors, fans [MW] 0.051total [MW] 0.141efficienciesenergetic efficiency, HHV [–] 79.4%energetic efficiency, LHV [–] 91.8%energetic efficiency, LHV, PE [–] 89.8%carbon yield [–] 80.2%

4.3.3 Exergy analysis

As described in section 3.5.1, the exergetic fuel and exergetic product can be defined intwo ways, firstly with the biomass as fuel and the torrefied wood as product (εI), andsecondly, with the biomass consumed in the reaction as fuel and the increase in specificexergy of the upgraded biofuel as product (εII). The exergy balance according to bothdefinitions is presented in Table 4.8. A detailed breakdown of the exergy losses and exergydestruction is given in Table 4.9.

93


The exergetic efficiency εI at 78.8% is relatively high, because most of the feedstock exergy“passes through” the torrefaction process and remains in the torrefied wood. However,the exergetic efficiency εII which takes into account the exergy that is actually convertedinto the desired product, namely the rise in specific chemical exergy of the biofuel, is only21.9%.

The biggest source of exergy destruction is, by far, the furnace, which contributes 42% ofthe total ED + EL, or 9% of the exergetic fuel of the plant. The exergy destruction fromthe torrefaction reactor amounts to 12% of total ED + EL, with an exergetic efficiency ofthe reactor of 70.3%. The drier accounts for 16% of the total ED + EL, and the exergyloss from drier exhaust gas for 15% of total ED + EL .

Given that most of the exergy destruction takes place because of drying and combustion,the focus for improving the torrefaction process should centre on the overall plant designrather than the torrefaction reaction itself. To improve the efficiency of the process, moreefficient drying technologies, such as SSD, should be investigated. Not only would SSDreduce the exergy destruction in the drier itself, but also decrease the amount of woodburned in the furnace. Integration of SSD in the torrefaction process, however, is not asstraightforward as for the wood pelletizing process. The hot combustion gas used for re-heating the torrefaction gas has a temperature of over 300°C at the heat exchanger outlet(21). In the presented process design, its thermal energy is used for drying the feedstock.If SSD is integrated in the process, this thermal energy would have to be utilized elsewherein the plant, which may not be feasible.

Table 4.8: Exergy balance for the torrefaction plant TOR-1.0.

εI εII

exergetic fuel [kWex] 13668 3713biomass [kWex] 13519 3563electricity [kWex] 145 145air [kWex] 4 4

exergetic product [kWex] 10771 814exergy losses [kWex] 476 476exergy destruction [kWex] 2422 2422exergetic efficiency [–] 78.80% 21.91%


The investment costs are summarized in Table 4.10. Information on the applied costfunctions and a detailed equipment list for TOR-1.0-m is provided in section B.2.2. Abreakdown of the production cost of the torrefied pellets is given in Table 4.11.

The specific investment costs amount to 188–400 €/kWbiofuel. According to RWE Innogy,the investment for their demonstration plant with a capacity of 60 kt/a of torrefied pelletsis 15 M€ [139], which equates to 306 €/kW.6 This is in good agreement with the valuescalculated in this work. Other study estimates, however, report considerably lower costs.

6Assuming a capacity factor of 80%.

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4.3 Torrefaction

Table 4.9: Exergy losses and exergy destruction for the torrefaction plant TOR-1.

[kW] % of ED + EL

exergy losses 476drier exhaust gas 425 15%ash 51 2%exergy destruction 2422torrefaction reactor 344 12%drier 473 16%furnace 1220 42%compressors, fans 13 0%gas mix 192 7%heat exchanger 31 1%cooler 59 2%pellet press 90 3%

Bergman et al. estimate the investment for a torrefaction plant with a capacity of 60kt/a to be in the range of 121–175 €/kW, without pelletizing [114].7 With pelletizing, thespecific investment costs for a 56 kt/a plant are estimated at 185 €/kW[117]. Verhoeff etal. report an investment of 17.1 M€ for a plant with a production capacity of 93 kt/a oftorrefied pellets [122], resulting in 193 €/kW.

The production costs of torrefied pellets from SR result in 9.0–11.3 €/GJ, or 186–233€/t. The feedstock costs including transport and storage contribute 60–80% of the totalproduction costs. The carrying charges contribute 12–20%. Compared to wood pelletsfrom the base case scenario WP-1.0, the pellet production costs are 0.5–0.8 €/GJ higher,which means a cost increase of 6–8%. The higher production cost is mostly due to thehigher investment cost of the torrefaction plant compared to the wood pelletizing plant.

In the case of overseas production, the higher density of the torrefied pellets leads to areduction in shipping cost by 0.35 €/GJ compared to wood pellets. The total costs, asdelivered to a port in Europe, are 0–0.3 €/GJ higher due to the more expensive upgradingprocess.

Similar production costs for the torrefied pellets are reported in the literature. Bergmanet al. estimate the production costs without feedstock costs as 55 €/t for a plant with acapacity of 230 kt/a of torrefied wood [114]. In this work, the production costs withoutfeedstock costs for TOR-1.0-m and TOR-1.0-l are 46–77 €/t.8 In contrast to this work,Bergman et al. conclude that, compared to conventional wood pellets, the production oftorrefied pellets from fresh wood is 0.9 €/GJ cheaper [117].

Studies which analyze the supply chain logistics with more detail reveal bigger cost savingsfor torrefaction compared to conventional wood pellets. Verhoeff et al. conclude that theproduction costs (without transport) are 0.2–0.3 €/GJ higher for torrefied pellets than forconventional wood pellets, and that these additional cost are more than compensated bysavings in shipping and storage of about 0.8 €/GJ [122].9

7Adjusted to €2010.8Without pellet transport and including the cost of wood used as auxiliary fuel at the torrefaction plant.9The pellets are produced in South Africa and shipped to Europe in this scenario.

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Table 4.10: Investment costs for the torrefaction plants processing short rotation wood.

TOR-1.0-s TOR-1.0-m TOR-1.0-l[M€] [M€] [M€]

torrefaction reactor 0.65 2.50 11.76drier 0.64 1.64 4.20furnace 0.21 0.61 1.77heat exchanger 0.15 0.35 0.93torrefaction gas compressor 0.14 0.52 1.77exhaust gas compressor 0.04 0.17 0.59air fan 0.00 0.00 0.01torrefied wood cooler 0.06 0.13 0.24pellet press 0.42 1.26 5.85pellets storage & handling 0.21 0.60 0.69total CBM 2.53 7.78 27.81offsite cost 0.30 0.88 2.55fees & contingencies 0.38 1.17 4.17start-up 0.15 0.50 2.00working capital 0.53 2.25 10.75AFUDC 0.32 0.98 3.45residual value (NPV) -0.10 -0.47 -2.31TCI 4.11 13.09 48.43

Table 4.11: Plant capacities, specific investment costs and levelized production costs oftorrefied pellets from short rotation (SR) and from forest residues (FR).

case TOR-1.0- s-SR m-SR l-SR s-FR m-FR l-FRfeedstock energy [MWHHV] 12.8 64.2 320.9 12.8 64.2 320.9biocoal energy [MWHHV] 10.3 51.5 257.7 10.3 51.5 257.7feedstock mass [ktar/a] 33.1 165.7 828.4 33.1 165.7 828.4biocoal mass [ktar/a] 12.6 63.0 315.2 12.6 63.0 315.2specific TCI [€/kWbiofuel] 399 254 188 381 251 170carrying charges [€/GJHHV] 2.23 1.42 1.05 2.13 1.40 0.95labour [€/GJHHV] 1.69 0.64 0.25 1.69 0.64 0.25electricity [€/GJHHV] 0.40 0.40 0.40 0.40 0.40 0.40O&M, material [€/GJHHV] 0.32 0.20 0.16 0.32 0.23 0.16other operating cost [€/GJHHV] 0.02 0.02 0.02 0.02 0.02 0.02biomass [€/GJHHV] 4.98 4.98 4.98 2.15 2.15 2.15biomass transport [€/GJHHV] 0.29 0.49 0.77 0.29 0.48 0.74biomass storage [€/GJHHV] 1.37 1.37 1.37 0.20 0.20 0.20pellets transport [€/GJHHV] 1.69 1.69 1.69biofuel cost [€/GJHHV] 11.31 9.53 9.00 8.90 7.21 6.57

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4.4 Anaerobic digestion


Two process designs involving anaerobic digestion are analyzed: the integrated generationof biogas and solid biofuel pellets, and the upgrading of the biogas to biomethane forinjection into the natural gas grid. Grass silage from conservation gras land is used as thefeedstock in both cases. Due to the low availability of the feedstock, only the small plantcapacity (denoted -s) with an input of 2.0 t/h of dry matter is considered.

The methane yield from the whole-crop digestion of grass silage is very low. Since themethane yield from grass varies widely depending on the cropping regime, a second, morefavourable simulation case with a higher methane yield is conducted. It should be notedthat such a high yield for grass from conservation grassland is not achievable with currentdigestion technologies. Rather, this case presents an upper bound in order to analyze thesensitivity of the energetic and economic performance in relation to the biogas yield.

The simulation case with solid biofuel pellet production is referred to as ADP-3.0, thecases with biomethane production with low and high yield as ADM-3.0 and ADM-3.1,respectively. Aspen Plus simulation models originally developed by Plaetrich [328] wererefined and adjusted to the methane yields, feedstock composition and general assumptionsused in this work.

4.4.1 Design and simulation model of anaerobic digestion with biomethaneproduction

The biomethane production process comprises anaerobic digestion and biogas condition-ing. The flowsheet of the analyzed design is shown in Figure 4.4. Biomass is mixed withwater recycled from the fermentation residues in the ratio 1:1 (K1) to form a slurry andis then fed into the digester. The slurry at the digester inlet has a solid matter content of15%. The digestate is separated into a liquid (10) and a thickened fraction with a solidmatter content of 30% (9) using a decanter. The liquid phase is modelled as pure water inthe simulation, with any solid matter passing into the liquid being neglected. The biogas,which is saturated with water vapour, is first cooled to 5°C (K3) in order to condenseand separate this water. It is then compressed to 8 bar in two steps (K5, K7). Thermalenergy is recovered by cooling the gas with water (K6, K8), which is then used to heatthe digester to its operating temperature of 39°C. After H2S and CO2 have been removedwith pressurized water scrubbing (K9), the biomethane is again cooled to 3°C (K10) toremove water (K11), and finally compressed to 16 bar for injection into the natural gasgrid. The thickened digestate is fed into a residue storage tank with the capacity to holdit for 100 days, in order to bridge the winter season when it cannot be spread on fields[221]. Data of the flow streams for ADM-3.0 is presented in section B.3.3. Details of thedigester design and heat loss are presented in section B.3.4.

4.4.2 Design and simulation model of anaerobic digestion with biofuel pelletproduction

The process design is based on the so-called integrated generation of solid fuel and biogas(IFBB) suggested by Richter et al. [53]. The flowsheet is shown in Figure 4.5. The

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9

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decanter

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flue gas

refrigerantelectricity

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Figure 4.4: Flowsheet of anaerobic digestion with biomethane production ADM-3.0.

10

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283132

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combustiblegas

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M

Figure 4.5: Flowsheet of anaerobic digestion with solid biofuel pellets ADP-3.0.

feedstock biomass is exposed to hydrothermal conditioning at 60°C (K1) by mixing withhot water (48) in the ratio 4.5:1. Press fluid and press cake are then separated with a screwpress (K2). 28% of the dry matter ends up in the press fluid. The dry matter content ofthe press fluid and press cake is 1.6% and 47%, respectively. Details on the mass flowsinto the press fluid and cake are given in section B.3.1. The press fluid is cooled to 48°Cand fed to the digester. The biogas is cooled to 3°C to condense the water (K16), andthen reheated to 70°C (K12) before combustion in a reciprocating engine. The exhaustgas at 454°C is used for drying the press cake. Since it does not provide enough thermalenergy to dry the press cake to the desired water content of 10%, additional hot gas fordrying (25) is produced by burning some of the produced press cake pellets (23). Thefermentation residues are separated into a liquid (10) and thickened digestate. Part ofthe liquid fraction is utilized for the hydrothermal conditioning after heating to 68°C withwaste heat from the engine (K6), press fluid (K5), and the drier exhaust (K7).

Digester heating is not required when operating at the design point because the press fluidenters the digester at a high temperature. In practice, additional heating capacity for coldweather may need to be considered.

Desulphurization is accomplished by sulphide precipitation in the digester. Though the

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chemical reactions are not considered in the simulation, the cost of iron oxide is accountedfor in the cost analysis.

Data concerning the flow streams is given in section B.3.3.

4.4.3 Simulation model of the anaerobic digestion reaction

A simple black box model for the anaerobic digestion of press fluid and whole crop silageis developed based on experimental data from [234] for the simulation cases ADP-3.0and ADM-3.0. Richter et al. conducted batch experiments on the anaerobic digestion ofhydrothermally conditioned grass silage from four different semi-natural grassland typestypical for German mountain areas. The average methane yield amounted to 445 l/kgoDMfor the press fluid and 226 l/kgoDM for whole crop digestion. The organic matter conversionwas 85% and 54% for the press fluid and for the whole crop silage, respectively [234]. Forsimulation case ADM-3.1, a very high methane yield is assumed, based on the yieldsachievable for grass from intensive cropping systems (see Table 2.10).

The simulation in Aspen Plus requires that all reaction products be defined. However, theelemental composition of the digestate is not given in [234], and the data on press fluid,press cake and feedstock is insufficient for the purpose of formulating elemental balancesfor C, O and H. Therefore, the model is partly based on assumptions, which are explainedin section B.3.2. Key results are summarized in Table 4.12.

The composition of the digestate and the heat of reaction are calculated by difference fromthe elemental and energy balances, respectively, and underlie a high degree of uncertainty.However, they are not central for the techno-economic evaluation of anaerobic digestionas an energy conversion technology. The digestate is a byproduct of the process whichis not analyzed further, and the heat of reaction makes only a minor contribution to theenergy balance.

Due to the different feedstock compositions, the calculated methane yield per kg of dryorganic matter is 5–13% lower for the simulation cases in this work compared to thosereported in [234]. The energy yield, however, is only 0.2–1.1 percentage points lower.

The assumed OLR in Table 4.12 are based on [233] for ADP-3.0 and ADM-3.0 and ondata for grass silage from [35] for ADM-3.1.

4.4.4 Energy balance

The main energy flows within the anaerobic digestion plant models are shown in Table 4.13.The residual is mostly heat losses and unused waste heat. In ADM-3.0, the energeticefficiency is only 41% due to the low digestibility of the feedstock. Indeed, 53% of thefeedstock energy is not converted by the anaerobic digestion and leaves the facility asdigestate. In ADM-3.1 with a more efficient conversion of the feedstock, the energeticefficiency increases to 74.5%. The main electricity consumers are the gas compressors,pressurized water scrubber, and digester stirrer.

In ADP-3.0, the energy ending up in the digestate is greatly reduced to 2% of the feedstockenergy. Instead, the hard-to-digest lignocellulosic material ends up in the press cake

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Table 4.12: Key data from the anaerobic digestion model.

ADP-3.0 ADM-3.0 ADM-3.1OLR [kgoDM/m3/d] 1.5 3.0 4.0press fluid / press cake separationDM content press cake [w%] 47.30% — —DM content liquor [w%] 1.64% — —mass yield press fluid (d.b.) [–] 27.7% — —energy yield press fluid (HHV) [–] 26.5% — —anaerobic digestionCH4 yield [l/kgoDM] 426.60 197.01 365.29CH4 concentration [mol%] 52.18% 46.1% 59.6%oDM conversion [–] 86.70% 53.37% 63.70%energy yield biogas (HHV) [–] 86.9% 43.2% 80.1%energy yields from feedstock (HHV)biogas [–] 23.0% 43.2% 80.1%press cake [–] 73.5% — —digestate [–] 2.0% 52.9% 16.6%heat of reaction [–] 1.5% 3.9% 3.3%

pellets. The main product is the press cake pellets, equivalent to 65% of the feedstockenergy. Surplus electricity produced by the CHP module is a byproduct with a minorcontribution equivalent to 6% of the feedstock energy. 36% of the electricity produced bythe CHP module is consumed by the plant, the main electricity consumers being the screwpress, pellet press, stirrers for hydrothermal conditioning and fermenter, and the decanter.Moreover, 12% of the produced press cake pellets (805 kW) are consumed internally toprovide thermal energy for the drier. The overall energetic efficiency is 71%.

Table 4.13: Energy balance of the anaerobic digestion plant models.

ADP-3.0 ADM-3.0 ADM-3.1inputbiomass [kWHHV] 9509 9509 9509electricity [kWel] — 387 606outputbiomethane [kWHHV] — 4064 7536press cake pellets [kWHHV] 6187 — —electricity [kWel] 532 — —digestate [kWHHV] 188 5029 1578losses1) [kW] 2601 802 1000energetic efficiency [–] 70.7% 41.1% 74.5%

1)heat losses, waste heat, exhaust gas, methane loss, unburned carbon in ash

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4.4.5 Carbon balance

Table 4.14 shows the carbon balance of the anaerobic digestion cases. When the upgradedbiofuel is utilized in a power station with CCS, the bulk of the carbon contained in thepress cake pellets and in the biomethane can be captured and stored. This amounts to65% for ADP-3.0 and 23–43% for the plants which produce biomethane.

For ADP-3.0, the remaining carbon is mostly released to the atmosphere in form of CO2from the combustion of biogas and pellets.

For the cases with biomethane production, 27–29% of the feedstock carbon is convertedto CO2 in the digester and released as offgas from the pressurized water scrubber. Itis assumed that the scrubber offgas is treated with a thermal oxidation process, whichconverts any remaining CH4 to CO2 [232]. In the analyzed plant design, the scrubberoffgas consisting mostly of CO2 is released to the atmosphere. However, it could easily becompressed for storage, if a connection to a CO2 transportation network were feasible.

49% and 27% of the feedstock carbon ends up in the digestate in cases ADM-3.0 and ADM-3.1, respectively. When the digestate is applied to agricultural land, some of the carbonis converted to soil organic matter, while the remainder forms CO2 by biodegradationand is released to the atmosphere. According to Fischer and Glaser [329], 35–45% of thedigestate carbon is typically converted to humus. However, only 4–14% of the carbon isstabilized in the soil at the end of the 100 years boundary for the GWP100, and can thusbe accounted for as sequestered carbon [330]. Carbon storage in the soil is not accountedfor in this work.

Table 4.14: Carbon balance of the anaerobic digestion plant models.

ADP-3.0 ADM-3.0 ADM-3.1press cake pellets 65.0% — —biomethane — 23.2% 43.0%scrubber offgas — 26.7% 28.9%digestate 2.7% 49.2% 27.2%drier exhaust 31.5% — —other 0.8% 0.9% 0.9%

4.4.6 GHG emissions

In addition to the supply chain emissions of the feedstock and the emissions embeddedin the electricity consumed, process related emissions also play a role for anaerobic diges-tion. These comprise fugitive methane emissions and emissions related to the subsequentspreading of digestate on agricultural land.

Fugitive methane emissions are caused by leakages, feedstock storage, digestate storageand spreading, and unburned methane in the CHP module exhaust [331]. According tothe IPCC Guidelines for National Greenhouse Gas Inventories, fugitive methane emissionsare generally between 0–10% of the produced methane. Where technical standards ensurethe prevention of unintentional emissions, they are negligible [278]. Clemens et al. meas-ured methane emissions from biogas plants in Germany and report 1.3–12 gCH4/kWhel,

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Table 4.15: GHG emissions [kgCO2,eq/GJbiofuel] from the anaerobic digestion plants.

ADP-3.0-s ADM-3.0-s ADM-3.1-sfeedstock, road transport 1.145 1.743 0.940auxiliary energy 1) -15.378 17.013 14.390fugitive CH4 emissions 1.586 4.502 4.502total -12.647 23.258 19.832

1) negative values denote electricity production and substitution of electricity from the German power mix.

corresponding to approximately 0.7–7% of the produced methane. Unburned methane inthe exhaust of the CHP module was found to be the biggest source [331]. Open storage ofthe digestate can cause up to 6.4 gCH4/kWel [331], but for new biogas plants gas-tight di-gestate tanks are mandatory in Germany under the Renewable Energy Sources Act [332].Jury et al. report methane losses for biomethane production at 1-4%.10 According toMeyer-Aurich et al., 0 to 2% fugitive methane emissions for state-of-the-art biogas plantsfulfilling current standards are a reasonable assumption [45]. It is assumed here thatfugitive methane emissions amount to 1% of the methane production.

The modelled plants produce 15–19 kt liquid residue and 2–23 kt solid digestate per year.Spreading of the digestate on agricultural land can have positive and negative effects onthe GHG balance. On one hand, GHG emissions arise from transport fuel consumption forthe distribution of digestate on the field, and through N2O emissions from biodegradationof the digestate. On the other hand, the digestate displaces mineral fertilizers and theupstream GHG emissions related to its production. The conversion of digestate organicmatter to soil carbon results in a net removal of CO2 from the atmosphere if the soil carboncontent at the end of the assessed time period (100 years for GWP100) has increased [330].

Fertilizer replacement and soil carbon storage are dependent on the composition of thedigestate, soil type, climate and cropping system [330]. Digestate from source separatedmunicipal solid waste results in an average net credit of 1.4 kgCO2,eq/GJbiogas based on datafrom [330], where credits for inorganic fertilizer replacement and soil carbon storage are al-most offset by N2O emissions from the digestate spreading.11 Vetter and Arnold calculate acredit of approximately 10 kgCO2,eq/GJbiogas for biomethane from energy crops, comprisingfertilizer replacement and humus reproduction [333]. Scholz et al. report GHG emissionsfrom digestate handling and spreading of 40 kgCO2,eq/MWhel (4 kgCO2,eq/GJbiogas) [229].In this work, it is assumed that the positive and negative effects of the digestate spreadingoffset each other equally. Thus no GHG emissions are assigned to the digestate.

The supply chain GHG emissions of biomethane and press cake pellets are shown inTable 4.15. The major cause of GHG emissions is the electricity consumption of the ADMplants, contributing more than 70% of the total. The net production of electricity inADP-3.0, and the replacement of electricity from the German power mix, leads to netnegative emissions for this case.

10This includes the methane slip in the purification process, which is assumed here to be oxidized in thescrubber offgas treatment process.

11Replacement of inorganic fertilizer: 20–28 kgCO2,eq/tbiomass,FM, soil carbon sequestration: 6.6–45 kgCO2,eq/tbiomass,FM, N2O emissions 33–60 kgCO2,eq/tbiomass,FM, methane production: 60Nm3/tbiomass,FM [330].

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The exergy balance for the three plant models with anaerobic digestion is presented inTable 4.16. As described in section 3.5.1, the exergetic fuel can be defined in two ways:firstly with the biomass as the fuel (εI), and secondly, with the biomass minus the digestateand liquid residue as the fuel (εIII). Both efficiencies are given in Table 4.16, the valuesfor exergetic fuel and exergy losses are related to the definition εI. A detailed breakdownof exergy losses and exergy destruction is given in Table 4.17.

Table 4.16: Exergy balance of the anaerobic digestion plant models (according to defini-tion of εI), and exergetic efficiencies εI and εIII.

ADP-3.0 ADM-3.0 ADM-3.1exergetic fuel [kWex] 10133 10514 10733

biomass [kWex] 10127 10127 10127electricity [kWex] — 387 606air [kWex] 6 — —

exergetic product [kWex] 7080 3826 7093biomethane [kWex] — 3826 7093biofuel pellets [kWex] 6549 — —electricity [kWex] 532 — —

exergy losses [kWex] 576 5496 2163exergy destruction [kWex] 2476 1192 1478exergetic efficiency εI [–] 69.9% 36.4% 66.1%exergetic efficiency εIII [–] 71.7% 73.7% 80.7%

With the efficiency definition εI, the exergetic efficiencies of ADM-3.0 and ADM-3.1 are36% and 66%, respectively. This low efficiency is due to the high amount of exergy thatremains in the digestate, 50% of the plant fuel for ADM-3.0 and 18% for ADM-3.1. Withthe efficiency definition εIII, the exergetic efficiencies are 74% and 81%. For ADP-3.0, thedifference between the two efficiency definitions is much smaller, at 70% according to εIand 72% according to εIII.

For the plants with biomethane production, the exergy loss associated with the digestateaccounts for 52–79% of the total ED + EL. Other significant sources of exergy destructionand loss are the digester (15–33%), the exergy loss associated with the scrubber offgas(3–6%), and the exergy destruction due to the gas treatment (2–7%).

For ADP-3.0, the main sources of ED + EL are the exergy destruction in CHP module(28%), drier (15%), digester (13%), and furnace (12%), and the exergy loss associated withthe drier exhaust (10%). The exergy loss associated with the digestate is much smallerthan that of the ADM cases at 7% of ED + EL.

The exergetic efficiency εI of the digester is 80% for ADP-3.0 and 86% for the plants withbiomethane production.

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Table 4.17: Exergy losses and exergy destruction for the anaerobic digestion plant models(according to definition of εI).

ADP-3.0 ADM-3.0 ADM-3.1

exergy losses [kWex] 576 5496 2163

digestate [kWex] 213 5287 1906

drier exhaust gas [kWex] 291 — —scrubber offgas [kWex] — 174 216

other [kWex] 71 35 41

exergy destruction [kWex] 2476 1192 1478

digester [kWex] 408 996 1200

digester heating [kWex] — 7 10

drier [kWex] 446 — —furnace [kWex] 366 — —liquor / press cake separation [kWex] 113 — —reciprocating engine [kWex] 845 — —heat exchangers [kWex] 40 — —gas treatment [kWex] 10 163 239

pellet press [kWex] 80 — —decanter [kWex] 56 18 18

other [kWex] 113 8 11


The investment cost for the three plant models with anaerobic digestion are summarizedin Table 4.18. A detailed equipment list for ADP-3.0-s and ADM-3.0-s and informationon the cost functions and assumptions used can be found in section B.3.5. A breakdownof the production costs for biomethane and press cake pellets is given in Table 4.19.

Table 4.19: Specific investment and levelized biofuel production costs for the anaerobicdigestion plants.

ADP-3.0-s ADM-3.0-s ADM-3.1-sfeedstock [kt/a] 46.720 46.720 46.720biofuel pellets [kt/a] 9.963 — —biomethane [kt/a] — 1.846 3.424specific TCI [€/kWbiofuel] 1221 2405 1221carrying charges [€/GJ] 6.81 13.41 6.81labour [€/GJ] 2.57 1.47 0.79electricity [€/GJ] −1.91 2.73 2.31O&M, materials [€/GJ] 1.18 1.26 0.63other operating cost 1) [€/GJ] 0.11 0.40 0.40feedstock (at farm gate) [€/GJ] 6.85 10.42 5.62feedstock transport [€/GJ] 0.84 1.27 0.69feedstock storage [€/GJ] 3.79 5.77 3.11total biofuel cost [€/GJ] 20.23 36.73 20.36

1) including cost of biomethane feed-in in the ADM cases

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Table 4.18: Investment costs for the anaerobic digestion plants.

ADP-3.0-s ADM-3.0-s ADM-3.1-s

[M€] [M€] [M€]digester 1.63 2.92 2.24fluid/cake separation 0.78 0.00 0.00product upgrading 1.42 2.23 2.62digestate processing 0.27 1.13 0.96CHP module 0.45 — —biomass sizing & sorting 0.16 0.16 0.16total CBM 4.72 6.44 5.98offsite cost 0.86 0.86 0.86fees & contingencies 0.71 0.97 0.90start-up 0.20 0.24 0.24working capital 0.53 0.50 0.53AFUDC 0.63 0.83 0.77residual value (NPV) −0.10 −0.07 −0.08TCI 7.55 9.77 9.20

It is assumed that the liquid residue and digestate are used as fertilizer, and that the costof spreading is offset by the avoided costs of not using mineral fertilizer.

The TCI of ADP-3.0 is approximately 20% lower than that of the biomethane plants. Themajor investment items are the digester (34–38% of CBM) and the product upgradingequipment (28–40%) for all three plants. The product upgrading comprises a scrubber,compression and drying for the biomethane plants and drying and pelletizing of the presscake for ADP-3.0. Digestate processing, comprising the decanter and the residue storagetank, contributes 16–18% of the TCI for ADM-3.0 and ADM-3.1. For ADP-3.0, this isreduced to 6% because of the lower amount of digestate. The hydrothermal treatment(including heat recovery) and subsequent separation of fluid and press cake is a significantcost factor for ADP-3.0 with 17% of CBM.

The specific investment cost is approximately 1200 €/kWbiofuel for cases ADP-3.0 andADM-3.1. For ADM-3.0 they are almost twice as high due to the low biomethane yield.That low yield also leads to a very high biomethane production cost of 35.9 €/GJ, 80%higher than that of ADM-3.1 at 19.9 €/GJ. The production cost of the press cake pelletsin ADP-3.0 is 19.7 €/GJ.

The feedstock costs including transport and storage contribute 45–51% of the total pro-duction cost for all plants. The carrying charges contribute 31–37%. Revenues fromelectricity sales cover 9% of the annual costs in ADP-3.0. For the biomethane productionplants, purchased electricity represents 8–12% of the biomethane production cost.

The production costs of biomethane from maize reported in literature centre around 20–25€/GJ[218, 334]. Maize silage has a similar methane yield to that assumed in ADM-3.1,and the resulting cost of 19.9 €/GJHHV in ADM-3.1 is plausible.12

12Feedstock costs as delivered including ensiling are 34 €/t (w.b.) for the grass silage in this work. Typicalvalues reported in literature are 25–35 €/t for maize silage and 38–45 €/t for grass silage [35, 221].

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The production cost of ADP and ADM cannot be directly compared because they cre-ate different products. However, it is clearly not economically sensible to manufacturesolid biofuel pellets when the more valuable biomethane could be produced at a similarcost. Thus, ADP only seems attractive for hard-to-digest feedstocks with a low methaneyield. The relative merits of solid biofuel and biomethane production are discussed insection 4.6.6, taking into account their subsequent utilization in thermal power stations.The biomethane is used in a gas-fired combined-cycle plant, whereas the solid biofuelpellets are combusted in a coal-fired power station.

4.5 Hydrothermal carbonization

The design of an efficient HTC process includes the choice of feedstock, the degree ofcarbonization and other operating parameters, and the plant layout. These general issuesare discussed in sections 4.5.1 and 4.5.2, before the selected simulation model and itsresults are presented.

4.5.1 General considerations for HTC as a pretreatment for combustion andgasification

From a thermodynamic point of view, pretreating biomass with HTC makes sense wherethe overall efficiency of HTC plus the subsequent combustion or gasification of the biocoalis higher than that of combustion or gasification of the raw biomass. This section analyzesthe influence of the degree of carbonization and the feedstock moisture content on theefficiency of combustion and gasification.

4.5.1.1 Influence of the degree of carbonization

As discussed in section 2.2.3, there is a trade-off between the higher heating value andthe energy yield. This raises the question as to the optimal degree of carbonization. Theminimum required degree of carbonization is defined by the desired change in hydrophobi-city and the loss of the fibrous nature. But should the biomass be carbonized further?For combustion and gasification applications, the overall conversion chain efficiency frombiomass to final product must be considered to decide this matter. Figure 4.6 shows theinfluence of the degree of carbonization of wood on the exergetic efficiency of combustionand gasification, and on the respective conversion chain efficiencies factoring in the energyyield of the HTC.13 Since this analysis focusses on the influence of the dry matter conver-sion, both wood and biomass are assumed to have a water content of 10%. The influenceof the water content is discussed in the next section.

Compared to the direct use of wood with an HHV of 19.5 MJ/kg, employing biocoal withan HHV of 28.7 MJ/kg increases the exergetic efficiency of combustion and gasification by13The oxygen-blown gasification is calculated under chemical equilibrium at 1550°C and 30 bar. The air

ratio of the adiabatic combustion is 1.2. The water content of wood and biocoal is 10%. The HTCreaction is modelled without dissolved organic compounds. The auxiliary energy consumption of theHTC process is not included in these figures.

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50%

55%

60%

65%

70%

75%

80%

85%

90%

19.5 21.5 23.5 25.5 27.5 29.5

biocoal HHV [MJ/kg]

exer

getic

effi

cien

cy

gasification

gasification, conv. chain

combustion

combustion, conv. chain

Figure 4.6: Effect of the degree of carbonization, represented by the HHV of the biocoal,on the exergetic efficiency of combustion and gasification, and the respective conversionchains taking into account the exergy yield of HTC.

3.8 and 4.3 percentage points, respectively. If the losses from the HTC reaction are factoredin, however, the efficiencies decrease by 1.1–1.9 percentage points. From a thermodynamicpoint of view, it therefore seems unnecessary, or even detrimental, to carbonize the biomassbeyond the point where the desired changes in its structure have been achieved.

For gasification, the degree of carbonization also has an influence on the syngas composi-tion. The impact of this on the efficiency of the overall process in which the gasificationis embedded, such as electricity or synthetic fuel generation, needs to be analyzed. Thegasification of wood and biocoal is discussed in detail in chapter 5.

The above considerations do not take into account the auxiliary energy consumption ofthe HTC process itself. A higher degree of carbonization also means that there is moreexothermal heat recoverable for feedstock preheating and product drying. The efficiencyof the internal heat recovery and subsequent auxiliary energy demand is dependent on thedesign of the HTC plant and requires detailed flowsheet simulations. Modelling results fordifferent HTC plant designs are presented in sections 4.5.4 and 4.5.12.2.

4.5.1.2 Influence of the biomass moisture

Another important question is which types of biomass should be upgraded using HTC.Since the reaction takes place in water, HTC seems particularly well suited to wet feed-stocks. In a combustion facility without flue gas condensation, the energy needed toevaporate the fuel moisture is lost. For fuels with a water content of over 50%, combus-tion therefore becomes very inefficient. HTC, on the other hand, facilitates the dewateringof the biocoal to a water content of 30–50% irrespective of the raw biomass moisture.

Figure 4.7 shows the energetic efficiency of a boiler burning raw biomass, and of HTC withsubsequent combustion of the biocoal, for boiler exhaust gas temperatures of 160°C and

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60°C. The water content of the HTC biocoal is 40%. With a boiler exhaust gas temperatureof 160°C, pretreating the biomass with HTC makes the overall conversion chain moreefficient if the biomass water content was originally higher than approximately 55%.14

With an exhaust gas temperature of 60°C, HTC is not worthwhile from a thermodynamicperspective. Since energy can be recovered by flue gas condensation, the efficiency forthe combustion of the raw biomass does not decrease so strongly with increasing moisturecontent. For combustion facilities where low temperature heat can be utilized and flue gascondensation is feasible, it is therefore thermodynamically better to burn the raw biomassthan to carbonize it first.

Combustion facilities designed for biomass are often equipped with flue gas condensation.For example, all wood-fired heating and CHP plants in Denmark include flue gas condens-ation with an exhaust gas outlet temperature below 65°C [335]. Standard coal-fired powerstations, however, usually do not recover the enthalpy of condensation.

30%

40%

50%

60%

70%

80%

90%

100%

0% 10% 20% 30% 40% 50% 60% 70% 80%raw biomass water content (w.b.)

ener

getic

effi

cien

cy (H

HV

) b

biomass, T=160°C

biomass, T=60°C

HTC, T=160°C

HTC, T=60°C

Figure 4.7: Relation between biomass moisture and energetic efficiency (HHV) of a boilerwith exhaust gas temperatures T of 160°C and 60°C, for the combustion of raw biomassand for the conversion chain of HTC and subsequent combustion.

The pretreatment with HTC can potentially increase the energetic efficiency of biomasscombustion when applied to feedstocks with a water content well above 50%, in cases whereflue gas condensation is not an option. From a thermodynamic perspective, pretreatmentwith HTC only makes sense under these circumstances. However, technical considerationssuch as better storage properties, better grindability and a higher energy density for longdistance transport may also play a role and widen the field of applications where HTCmay be attractive.

The data presented in Figure 4.7 does not include the auxiliary energy demand of theHTC process and therefore represents an upper limit of the HTC efficiency. Detailed

14The exact value may vary with the energy yield of the particular HTC reaction. The values presentedhere refer to strong carbonization with a biocoal HHV of 29.3 MJ/kg and an energy yield of 83.2%.

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simulation studies of HTC and subsequent combustion for selected feedstocks are presentedin section 4.5.15. These studies do include the auxiliary energy demand.

4.5.2 Considerations for an industrial-scale plant design

One of the most important design decisions for an HTC plant is whether the reactorshould be operated in continuous or batch mode. An efficient heat recovery is mucheasier to realize with a continuous operation. In batch mode, the recovered steam requiresintermediate storage of and/or several reactors which operate in a time-staggered sequence.A typical HTC residence time of several hours leads to long idle periods if only two reactorare employed [199]. Technology developers pursuing batch concepts therefore envisagedesigns with 6–12 reactors [175].

A key drawback of continuous operation is that the biomass feedstock must be fed intothe reactor against a pressure of 10–40 bar. This is technically challenging. Bloeß [336]conducted a literature and patent review and concludes that specialized pumps and plugforming feeders are the most suitable systems, while lock hoppers are not suitable dueto the large amount of inert gas they require. Specialized piston pumps are already inoperation in HTC pilot plants but have the disadvantage that the biomass needs to bediluted with water to make the slurry pumpable. This will adversely affect the energeticefficiency, since the additional water needs to be preheated to the reaction temperature.One manufacturer suggested dewatering the biomass with a screw feeder after pressur-ization. Plug screw feeders will work with lower water content (approx. 30%) and arecommonly used for biomass compression in the pulp and paper industry for pressure dif-ferences of up to 10 bar. There is, however, almost no experience with the higher pressuresrequired in an HTC plant. Bloeß reports that several manufacturers indicated in personalcommunications that pressures of 20–25 bar are technically feasible, but have not beenpursued in the past due to lack of demand and high cost. With the current rise in demandfor such systems, for example for the pretreatment of wood for bioethanol production, sev-eral manufacturers have indicated development projects for pressurizing biomass to higherpressures.

Swanson et al. analyzed a wide variety of feeding systems with regard to their suitabilityfor pressurized biomass gasification. They identify the plug-screw feeder as the technologywith the best track-record for high-pressure biomass feeding [337, page 106]. Abrasion hasproved an issue for pumps in hydrothermal peat upgrading plants [200].

Without heat recovery, approximately 15–30% of the biocoal energy is required for pre-heating the biomass slurry. With complete heat recovery, on the other hand, the heatof reaction and the thermal energy recovered from the cooling of the products is suffi-cient for preheating the biomass and drying the product. The heat recovery scheme istherefore of great importance for the energetic efficiency of the process. The literatureon the hydrothermal upgrading of peat [154, 200, 201] includes flowsheet designs withvarious heat recovery schemes, which could well be applicable to modern HTC plants. Attemperatures above 100°C, the peat and product slurries caused problems with foulingand clogging [188]. Therefore, most peat upgrading processes employed only direct heattransfer at elevated temperatures [154, 200, 201]. A common scheme was a counterflowarrangement whereby the pressure of the product slurry is stage-wise reduced and the

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flash steam condenses in the feed slurry. However, the use of heat exchangers at hightemperatures was pursued in some plant designs. In a pilot plant using the DeLaval’swet-carbonization process, a specially designed heat exchanger applying centrifugal forcesto the slurry was employed for temperatures up to 200°C [154].

Other important design decisions include the type of drier, the provision of additionalprocess steam, and the design of the mechanical dewatering facility.

The flowsheet design of the HTC base case presented in this work most closely resemblesa peat upgrading pilot plant with continuous reactor operation and direct heat transferwhich was operated in Sweden from 1954 to 1964 [200]. The successful operation of thepeat processing pilot plant for 10 years demonstrated the general technical feasibility ofthe process design.

4.5.3 Simulation model of the HTC reaction

The HTC reactor is modelled as a black box based on measured yields and compositionsfrom a series of HTC experiments conducted by Stemann [23, 163, 185] with poplar woodand EFB at 220°C with a residence time of 4 h. Only experiments conducted in recirculatedprocess water are considered, since they better reflect the conditions found in an industrial-scale process. The recirculation of process water was shown to increase the biocoal yieldand the HHV and improve the susceptibility to mechanical dewatering (see section 2.2.3.9).The experimental setup and key results of the experiments are described in [23, 163]. Firstscale-up trials indicate that results from laboratory batch experiments can be reproducedon a pilot plant level under a continuous operation [197].

The measured data comprises the elemental composition of the feedstock and biocoal, thebiocoal yield, the yields of the gaseous components CO2, CO, CH4 and H2, the TOC andthe yields of dissolved organic compounds identified by HPLC analysis. As discussed insection 2.2.3.5, part of the TOC is present as organic compounds which are not identified.It is therefore unknown how much of the hydrogen and oxygen that is removed from thefeedstock has reacted to water and how much has been converted to dissolved organiccompounds. For the purpose of simulation in Aspen Plus, assumptions on the hydrogenand oxygen conversion are required, which fulfil the material balances and which leave thecomposition of the dissolved compounds in a plausible range. It is assumed that the molarratio r of CO2 to H2O formed in the reaction is 0.163 for the HTC of wood.15 This resultsin a molar O/C ratio of 0.526 and a molar H/C ratio of 1.963 for the dissolved organiccompounds. For EFB, r is set at 0.282, O/C is 0.437 and H/C is 1.980.

In the simulation, acetic acid and formic acid are considered separately, while the residualof the dissolved organics is modelled based on its aggregated elemental composition. This isreferred to as TOMres (total organic matter residual). Mineral compounds are aggregatedas “ash”.

The fate of nitrogen and sulphur removed from the solid phase is unknown. H2S has notbeen measured in the experiments, and N and S compounds have not been detected in15Values for r reported in literature are mostly higher, in the range of 0.2 to 0.4 [153, 158]. However,

they are often calculated without considering the dissolved compounds, as if all carbon and hydrogenremoved from the feedstock react to CO2 and H2O. These values are therefore not directly comparable.

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the dissolved substances. It is assumed that the balance of S reacts to H2S (since H2S hasbeen reported as an HTC reaction product [187]), and that the balance of N ends up inthe dissolved compounds.

For MOW, no measured data was available. It is therefore assumed that the O/C ratioin the resulting biocoal, the gas phase composition, the composition of the TOMres andr are the same as that for wood as a feedstock. The resulting CO2 and H2O were thencalculated by difference from the carbon and oxygen balances.

Experiments with high ash biomass show that ash accumulates in the biocoal and a smallershare of the feedstock ash ends up in the liquid phase [159]. It is assumed that 0.03 kgash per kg dry PGW or MOW are dissolved while the rest remains in the biocoal.

To investigate the influence of the degree of carbonization on the efficiency and cost of anHTC plant, a case with weak carbonization (210°C, 3 h) and a case with strong carboniz-ation (230°C, 8 h) are also considered for PGW-70. Based on experimental data for beechwood at 200°C and 250°C, both at 4 h [158], the following relation between the O/C ratioof the biocoal and the carbonization temperature [°C] for two operating points (1, 2) wasestimated:(

O

C

)2

=(

O

C

)1

− 0.004231 (T2 − T1) (4.1)

It has to be pointed out that this relation can only be a rough estimate, because it wascalibrated using only two data points for beech wood and it is questionable whether it canbe legitimately transferred to other types of biomass.

Within certain limits, temperature T and residence time t are substitutable. The relationbetween temperature, residence time and degree of carbonization given by Equation 2.3 isapplied to define two additional simulation cases. Both have the same degree of carbon-ization as the base case HTC-3.00 but different T,t pairings, namely 210°C at 8.4 h and230°C at 2 h.

Table 4.20 shows the resulting biocoal yields and compositions in the various simulationcases. The composition of the feedstocks is given in Table 3.3. The yields refer to thefinal product of the HTC plant. In the drier, the composition changes slightly due tocondensation of inorganic and organic compounds on the biocoal. The composition ofthe final product therefore varies slightly in different simulation cases depending on thedegree of mechanical dewatering and the water to biocoal ratio in the slurry. Values forthe composition of the biocoal at the reactor outlet and the yields of all byproducts aregiven in section B.4.1.

The energy yield for HTC from EFB is 10 percentage points lower than that from wood,mostly due to a higher share of dissolved organic compounds. Since the simulation modelfor PGW, MOW and grass was derived from that of wood, there is a high uncertaintyregarding the energy yield. The biocoal from MOW has a considerably lower calorificvalue than that of the other feedstocks due to its high ash content.

Before entering the reactor, the biomass slurry has to be preheated. The minimum temper-ature is the reaction onset temperature of approximately 180°C, but in practice a higherslurry temperature is required to reach the final reaction temperature of 220°C. The tem-

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Table 4.20: Biocoal composition and yields in the simulation cases.

HTC-1.00 HTC-3.00 HTC-4.00 HTC-5.00 HTC-3.01 HTC-3.02

feedstock wood PGW, grass MOW EFB PGW PGW

T [°C] 220 220 220 220 210 230

t [h] 4 4 4 4 3 8

carbon [w%] 65.96% 62.50% 47.95% 64.32% 59.39% 67.66%

hydrogen [w%] 5.56% 4.70% 3.61% 6.92% 4.89% 4.38%

nitrogen [w%] 0.11% 1.97% 1.51% 0.84% 1.85% 2.15%

sulphur [w%] 0.05% 0.18% 0.14% 0.07% 0.17% 0.20%

oxygen [w%] 26.69% 25.33% 19.44% 24.21% 28.65% 19.78%

ash [w%] 1.64% 5.33% 27.36% 3.64% 5.05% 5.83%

HHV (d.b.) [MJ/kg] 26.784 24.607 18.393 28.017 23.409 26.601

mass yield [–] 63.41% 57.96% 64.06% 58.66% 61.49% 52.91%

energy yield [–] 86.88% 83.32% 83.07% 77.05% 84.10% 82.23%

perature to which the biomass slurry has to be preheated depends on the HTC heat ofreaction which contributes to the further heating of the slurry once inside the reactor.

In the simulations, 3–7% of the feedstock energy is converted to thermal energy in theHTC reactor, corresponding to 0.65–1.27 MJ/kg feedstock (daf). This is within the sameorder of magnitude as the measured heat of reaction from the HTC of wood and cellulose,which was found to lie in the range of 0.76–1.07 MJ/kg feedstock (daf) [191].

The process design should take into account that a significant amount of the heat ofreaction is consumed by evaporating water. Since the gas phase is saturated with watervapour, the amount of water evaporated depends on the amount of gaseous byproductsformed in the reaction and on the pressure and temperature inside the reactor. In thebase case HTC-3.00, 82% of the heat of reaction is consumed to evaporate water, whichleaves the reactor with the gaseous byproducts. With a reactor operating temperature of220°C and a pressure of 25 bar in HTC-3.00, the molar fraction of water vapour in the gasphase is 92%. For each kg of CO2 formed, 5.6 kg of water is evaporated. When the reactorpressure is increased to 30 bar, this reduces to 1.65 kg water per kg CO2. Less evaporationmeans that a greater share of the heat of reaction is available for heating the slurry tothe final operating temperature of 220°C. Thus, the required temperature to which thebiomass slurry needs to be preheated in order to fulfill the energy balance of the reactordecreases with increasing pressure. For the reaction at 25 bar, the biomass slurry has tobe preheated to 214°C. For the reaction at 30 bar, 193°C is sufficient. It may thereforebe advantageous to operate the reactor at a pressure well above the saturation pressure,in order to minimize the energy loss by evaporating water. Although the HTC reaction isitself exothermic, it is possible that the thermal energy consumed by evaporation exceedsthe heat of reaction, in which case heating of the reactor would be required.

4.5.4 Design and simulation model of the HTC base case

Figure 4.8 shows the flowsheet of an industrial-scale HTC plant with a continuous reactorand a feedstock processing capacity of 6667 kg/h park and gardening waste (FM).

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The biomass is first mixed with recirculated process water to create a pumpable slurry witha dry matter content of 15% (K1). It is then pressurized and preheated in several stages.Heat exchangers are only employed at temperatures below 100°C, because decomposingbiomass and condensing tar-like substances from the product would likely cause foulingand clogging at higher temperatures. Above 100°C, the slurry is preheated by mixing withsteam recovered at different temperatures during the cooling and de-pressurizing of thebiocoal slurry. The pressure levels of the first three de-pressurization stages (K17, K18,K19) are chosen so that the total amount of steam recovered in each step can be utilized inthe respective mixing pre-heater.16 The heat recovery scheme is adapted from pilot plantsfor the hydrothermal treatment of peat [154, 200, 201]. The additional steam requiredto reach the reaction temperature is produced by a boiler (K25) fired with dried biocoal.The HTC reactor is operated at 220°C and 25 bar. The biocoal slurry is mechanicallydewatered (K21) to a water content of 40%. Water from the filter press is recovered toprepare the biomass slurry, the remainder is cooled to 40°C and subjected to aerobic wastewater treatment (94). Condensate collected at various stages of the process (36) is alsosubjected to the aerobic waste water treatment. The biocoal is then dried to a watercontent of 10% (w.b.) using low temperature drying. To this end, the drying air is heatedwith steam (K32) at 100°C recovered from the biocoal slurry depressurization process,and by mixing with the boiler exhaust gas (K39). It is assumed that 30% of the TOMresare lost as VOC emissions from the drier (99). Finally, the dried biocoal is pelletized.The gaseous byproducts (87) from the HTC reactor consist mostly of water vapour andCO2. Because these byproducts also contain some CO, they need to be further oxidized.They are therefore cooled to condense the water and then co-combusted in the boiler.Heat from the gaseous byproducts is recovered to produce steam at 11 bar (K41) and topreheat the boiler feedwater (K26). Data on the material and energy flows can be foundin section B.4.2.

biocoal

condensate

ashreactor

pelletpress

flue gas

filter pressair

76

50

14

47

19 20 21 22 23

6

15

13

96

54

53

1639

37 4

10

7

5 67

8011 86

12

87

82

81

56

62

92

41

97

61

28

63

78

8564

70

59

70

72

100

60

29

90

7873

63

95

95

93

93

9

32

34

46

91

55

8340

8

57

44

36

5868+99

30

42

K2 K3 K4 K5

K28

K24

K45

K6 K7 K8 K9 K10 K11

K16 K15

K29

K19

K31

K21 K37 K44K46K22

K20

K18

K41

K40

K36

K27

K12

K39

K38K32

K26

K25

K42

K43

K13

K14

K17

17 71

air

drier

A

A

BC

C

D

E

E

D

B

biomass

wastewater

aeration

discharge

2794

1 K1

K33

3

2W1

W14

W2 W4

W9W7

W5

W11

W8

W3

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

Figure 4.8: Flowsheet of the HTC base case HTC-3.00.

16To avoid boiling of the slurry in the mixing preheaters, the maximum slurry outlet temperature is setat 5°C below its saturation temperature.

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The flowsheet design is exactly the same for the simulation cases HTC-1.00, HTC-2.00and HTC-4.00. For HTC-5.00 with EFB, it is slightly modified: Surplus fibres and shellsfrom the palm oil mill are used as fuel for the boiler instead of biocoal, and the boiler fluegas, which consequently has a high water content, is not utilized as a drying agent. Thereactor pressure is 30 bar. A flowsheet and data are given in section B.4.3. All simulationcases are based on 2000 kg/h dry matter biomass input, the fresh matter input varies inrelation to the feedstock water content.


The energy balance and efficiencies for the HTC base design, using various feedstocks, aregiven in Table 4.21.

The process efficiency on an HHV basis lies between 68–78%, the higher the ash andmoisture content of the feedstock, the lower the efficiency. On an LHV basis, however,the efficiency is highest for the feedstocks with the highest moisture content. Indeed,it is greater than 100% for the feedstocks with a 70% water content, namely PGW-70and MOW. Due to the mechanical dewatering of the biocoal, water is removed from theprocess in liquid state. Thus energy which would be required for evaporating the feedstockmoisture is saved, and more useful energy can be gained from burning the biocoal thanfrom burning the raw biomass. The LHV-based efficiency, taking into account the primaryenergy demand for the consumed electricity ηLHV-PE, is 5–15 percentage points lower thanthe simple LHV-based efficiency, but still greater than 100% for the feedstocks with 70%water content.

The electricity consumption amounts to 40–84 kJel/MJbiocoal. The largest electricity con-sumer by far is the aerobic waste water treatment, contributing 48–76% of the total electri-city demand of the HTC plant. Other important consumers are the pellet press (9–20%),slurry pumps (8–16%), fans for the drier and coolers (6–13%), and the dewatering press(2–4%). For details see Tables B.18 and B.20.

Moreover, 7–14% of the produced biocoal is used to fuel the boiler. 7–18% of the bio-mass energy is lost as chemically bound energy in dissolved organic compounds in thewaste water stream, and 12–21% of the biomass energy is lost because of heat losses andunrecoverable thermal energy in the exhaust gas and waste water streams.


Table 4.22 shows the carbon yields of the base case simulation for all considered feedstocks.The net carbon yield in the final product lies between 72% and 82% and is 5–12 percentagepoints lower than that of the HTC reaction itself, because part of the biocoal is burnedon-site. Simulation case HTC-5.00 has the lowest carbon yield due to the significantlyhigher proportion of the original carbon forming dissolved byproducts. Carbon in thedissolved byproducts (5–15%) and CO2 from the biocoal combustion (11–18%) are thebiggest losses. Less significant losses include the organic compounds evaporated in thedrier (0.7–1.8%) and unburned carbon in the boiler ash (0.2%).

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Table 4.21: Energy balance and efficiencies for the HTC base design for various feedstocks.

case HTC-1.00 HTC-2.00 HTC-3.00 HTC-4.00 HTC-5.00

feedstock wood PGW-50 PGW-70 MOW EFB

input

biomass [kWHHV] 10860 9509 9509 7879 11851

auxiliary boiler fuel [kWHHV] — — — — 793

electricity [kWel] 357 355 413 380 765

output

biocoal (net) [kWHHV] 8759 7404 7200 5647 9131

loss

dissolved compounds [kWHHV] 792 761 904 761 2237

heat losses, waste streams [kWth] 1,666 1,698 1,817 1,852 2,041

efficiencies

HHV [–] 78.1% 75.1% 72.6% 68.4% 68.1%

LHV [–] 91.1% 90.1% 113.0% 118.7% 86.0%

LHV-PE [–] 85.5% 83.7% 101.4% 104.0% 76.3%

biocoal combusted [%] 7.2% 8.1% 9.1% 13.7% —

Table 4.22: Carbon yields of biocoal and byproducts for the HTC base design.

HTC-1.00 HTC-2.00 HTC-3.00 HTC-4.00 HTC-5.00biocoal 81.7% 79.3% 77.4% 73.5% 71.9%CO2 11.2% 13.1% 13.8% 17.6% 11.5%waste water 5.1% 5.6% 7.9% 7.7% 14.7%VOC from drier 1.7% 1.8% 0.7% 0.8% 1.6%unburned C in ash 0.2% 0.2% 0.2% 0.4% 0.2%

4.5.7 GHG emissions

The GHG emissions for selected cases are shown in Table 4.23 and details on all casesare given in Table B.41. The GHG emissions using SR wood are the lowest for HTC,but still about twice those of the respective wood pelletizing and torrefaction cases. ForHTC-3.00-m using PGW, the emissions are about three times as high as those using wooddue to higher auxiliary energy consumption and road transport of the PGW. Due to thelow energy density of waste and its low yield per area, road transport is the biggest GHGcontributor for HTC-3.00-m.

For HTC-5.00-m from EFB, the avoided CH4 and N2O emissions from EFB dumping faroutweigh the CO2 emissions from transport and auxiliary energy consumption, resultingin net negative emissions of -173 kg/GJ.

GHG emissions of the electricity consumed by the HTC plant in case HTC-5.00-m arebased on the Malaysian power mix. This value would apply if electricity was taken fromthe grid. If electricity is produced on-site by the combustion of shell and fibres, the valuecan be as low as zero.

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Table 4.23: GHG emissions [kgCO2,eq/GJbiocoal,HHV] from HTC biocoal production forselected cases.

HTC-1.00-m HTC-3.00-s HTC-3.00-m HTC-5.00-mfeedstock, road transport 0.754 3.728 10.913 0.000biofuel, road transport — — — 1.344biofuel, ship transport — — — 5.646auxiliary energy 7.300 10.256 10.256 20.712avoided CH4 + N2O — — — -200.619total 7.753 13.984 21.169 -172.916

4.5.8 Environmental considerations

The HTC plant produces three waste streams: the exhaust gas from the drier, ash from theboiler, and waste water. The exhaust gas contains the combustion products from biocoaland the gaseous byproducts of the HTC reaction and may contain SO2, particulates andVOC from the drying of the biocoal. VOC emissions are an issue with biomass driers[325], and it has to be assessed whether this is also the case for biocoal drying.

Another issue is the retaining of some of the mineral content in the biocoal. If the biocoal isco-combusted with fossil coal, these nutrients can, most likely, not be recovered and cycledback into nature. If HTC replaces practices like mulching with EFB or composting, wherethe nutrients are returned to the agricultural land, the nutrients exported with the biocoalmay need to be replaced by mineral fertilizer.

Sulphur emissions, waste water and the fate of nutrients are discussed in the following.

4.5.8.1 Sulphur emissions

The simulation model for the HTC reaction assumes that the net difference of sulphur inthe feedstock and the biocoal forms H2S in the HTC reaction. This leads to a concentrationof more than 3000 mg/m3 H2S in the reactor offgas. H2S is converted to SO2 in the utilityboiler of the HTC plant, where the reactor offgas is co-combusted with biocoal. This leadsto a concentration of 1000 mg/m3 SO2 in the exhaust gas of the boiler, corresponding to1.1 kg/h SO2 for the 11 MW plant. This however is a worst case assumption, since partof the sulphur is probably present in the dissolved substances. Since H2S has not beenmeasured in the experiments on which the model is based, there is a high uncertaintyregarding the H2S content of the reactor offgas. The H2S concentration will probablybe strongly dependent on the sulphur content of the feedstock. Maximum permissibleSO2 emissions for combustion applications are dependent on plant capacity and fuel. Forcombustion applications, SO2 reduction measures are usually not required if low sulphurbiomass such as untreated wood is used as a fuel [28]. Sulphur emissions from HTC andthe necessity of remedial treatment require further investigation.

In the simulation and economic analysis presented here, desulphurization of the reactoroffgas is not considered. If desulphurization is required, H2S can be removed from thereactor offgas with technologies employed in anaerobic digestion plants, such as adsorptionwith activated carbon.

116


4.5.8.2 Waste water

For each kg of biocoal, the HTC plant produces 0.8–3.5 kg of waste water from dewa-tering the biocoal and 1.2–1.9 kg condensate from the flash steam. The condensate maycontain volatile organic compounds such as acetic acid, formic acid, phenol and light hy-drocarbons. A phenol concentration of 70 mg/l was measured in the condensate of steamobtained from the flash tanks in a plant for upgrading peat [200]. The waste water fromdewatering the biocoal contains a wide range of organic substances and dissolved min-eral compounds originating from the biomass. TOC of up to 33 mg/l were measured inlaboratory experiments (see section 2.2.3.5). Both the waste water from the biocoal dewa-tering and the condensate may include persistent organic pollutants and substances whichare hazardous to waterways, and therefore require treatment. 25–47% of the total wastewater and condensate can be recycled for steam production in the HTC plant after suchtreatment. The amount of waste water that requires discharging depends strongly on themoisture content of the feedstock.

In this work, waste water treatment is undertaken in oxidation ponds. The electricityconsumption of the waste water aeration is estimated with a simple model (section 3.3.2.9)and underlies a high uncertainty due to the incomplete knowledge about the amount andnature of the dissolved substances. Optimizing the treatment technology may reduce theelectricity demand. On the other hand, persistent pollutants may increase both the energydemand and the cost of the treatment. The type of treatment is likely to have a significantimpact on the cost of the biocoal production and needs further investigation.

4.5.8.3 Nutrient recovery

The consequences of nutrients remaining in the biocoal are analyzed for the case of EFB.

Some of the nutrients end up in the liquid byproducts. During aerobic treatment of thewaste water, the organic compounds are removed, yet the minerals remain in the water.By spreading the treated waste water on the plantation, nutrients are returned. However,it needs to be analyzed whether this is economically viable, whether the water containsharmful substances which prevent its application, and whether the nutrients are presentin a plant-available form.

In laboratory-scale experiments, 23% of the nitrogen and 72–88% of P, K, Ca and Mgended up in the HTC liquid byproducts [23].

Since both EFB and recovered nutrients from HTC waste water can replace mineral fer-tilizer, each can be assigned a value in terms of the avoided cost and GHG emissionsassociated with the displaced mineral fertilizer.

Based on cost and embedded GHG emissions of N, P and K fertilizers with the samenutrient value, the avoided GHG emissions of applying EFB to the plantation amount to26.9 kgCO2,eq/tDM of EFB and the monetary value is 4.4–11.7 €/tDM, or 1.5–4.1 €/tFM.52–68% of the monetary fertilizer value and 43% of the avoided GHG emissions are re-lated to the nutrients remaining in the HTC waste water and may therefore be partlyrecoverable. Details of these calculations are given in section B.4.4.

117


In addition to the value of the displaced fertilizer, applying EFB to the plantation mayimprove the structure and moisture retention of the soil, resulting in an increased FFByield. These effects would need to be assessed if HTC were to be compared with EFBmulching.


Similar to torrefaction, the exergetic efficiency of HTC can be defined in two ways. Theexergy balance for simulation case HTC-3.00 according to both definitions is presentedin Table 4.24. A detailed breakdown of exergy losses and exergy destruction is given inTable 4.25.

Table 4.24: Exergy balance of HTC-3.00.

εI εII

exergetic fuel [kWex] 10573 5274biomass [kWex] 10127 4827electricity [kWex] 413 413air [kWex] 33 33

exergetic product [kWex] 7490 2189exergy losses [kWex] 1108 1108exergy destruction [kWex] 1977 1977exergetic efficiency [–] 70.8% 41.5%

Table 4.25: Exergy losses and exergy destruction for HTC-3.00.

[kWex] % of ED + EL

exergy losses 1108drier exhaust gas 156 5%waste water 878 28%condensate 52 2%other 22 1%exergy destruction 1977HTC reactor 546 18%drier 173 6%boiler 483 16%slurry pumps 26 1%slurry preheaters (indirect) 63 2%slurry preheaters (mixing) 120 4%heat exchangers reactor gas 40 1%flash tanks 82 3%coolers 36 1%piping heat losses 16 1%filter press 43 1%milling and pelletizing 68 2%other 280 9%

118


As in torrefaction, most of the feedstock exergy remains in the solid matter and “passesthrough” the carbonization process, resulting in a high exergetic efficiency εI of 70.8%. Theexergetic efficiency εII, taking into account the exergy that is actually converted into thedesired product, is 41.5%. Because of the stronger conversion compared to torrefaction,εI is lower, but εII is almost twice as high as in TOR-1.0.

The main sources of exergy destruction and loss are the HTC reactor, the boiler, the drierand the dissolved organic compounds exported with the waste water, condensate and driereffluent. The losses of organic compounds amount to 28% of the total ED + EL or 8% ofthe exergetic fuel of the plant. It should be noted that there is a significant uncertaintyabout this value, since the behaviour of a continuous flow reactor with flash cooling mightbe different from the bench-scale experiments with slow cooling of the reaction products.The exergy destruction from the HTC reaction itself amounts to 18% of total ED + EL

and that of the boiler is 16%. The exergetic efficiency of the HTC reactor is 83%, that ofthe boiler 37%. The drier exergy destruction plus the exergy loss from the drier exhaustgas accounts to 11% of the total ED + EL. Detailed data on the exergetic efficiency andexergy destruction of all plant components can be found in section B.4.6.

For improving the efficiency of the HTC, the following potentials are identified based onthe exergy analysis.

In order to reduce the loss of chemical exergy with the waste water, either the amount ofdissolved organics in the reactor has to be reduced, or the exergy in the waste water hasto be recovered as useful products. Experiments with anaerobic degradation indicate thatbiogas production from HTC waste water offers some promise [220, 338]. Another approachis the separation and recovery of high value chemicals, such as phenol or acetic acid. Inthis case, reaction conditions may need to be modified to enhance the formation of thedesired compounds. Suppressing the formation of dissolved compounds and enhancing theformation of high value chemicals both require further research into the reaction chemistry.That task lies oputside the scope of this work. Recovering methane from the waste waterusing anaerobic digestion is discussed in section 4.5.13.

Exergy destruction in the boiler can be reduced by limiting the steam requirement. Thiscould be achieved by improving the preheating scheme or by decreasing the water tobiomass ratio in the slurry. The extent to which this is possible depends on the re-quirements of the feeding system. Some alternative configurations of the preheating areanalyzed in section 4.5.12 and the influence of the water to biomass ratio is discussed insection 4.5.12.1.

A more efficient drying system should also be investigated. To this end, superheated steamdrying with condensation of the exhaust gas moisture is discussed in section 4.5.12.


The investment cost for the plants processing PGW-70 are given in Table 4.26. Thetotal capital investment is 10 million € for HTC-3.00-s and 31 million € for HTC-3.00-m.The most expensive components are the HTC reactor (24–33% of CBM), the filter press(12%), and the biocoal drier (10–12%). There is a relatively high uncertainty about thecomponents which process the biocoal. Cost estimates based on several sources deviate

119


from each other by a factor of 1.8–2.8 for these components. Details about assumptionsand results for the cost estimates are provided in section B.4.5.

Table 4.26: Investment cost for HTC-3.00-s and HTC-3.00-m.

HTC-3.00-s HTC-3.00-m[M€] [M€]

HTC reactor 1.77 7.59slurry pumps 0.47 1.37flash tanks 0.53 1.70filter press 0.74 2.30biocoal drier (incl. HX) 0.78 1.93pellet press 0.34 1.00boiler 0.23 0.84heat exchangers 0.51 0.91coolers 0.35 0.73pumps and fans 0.04 0.07biomass slurry preparation 0.09 0.25biocoal storage & handling 0.17 0.48biomass sizing & sorting 0.16 0.44waste water treatment 0.34 0.88total CBM 6.52 20.50offsite costs 1.08 3.12fees & contingencies 0.98 3.08start-up 0.30 0.86working capital 0.43 1.04AFUDC 0.86 2.67residual value (NPV) -0.05 -0.10TCI 10.10 31.17

Plant capacities, specific investment and biocoal production costs for the HTC base designwith all feedstocks are given in Table 4.27.

The specific investment is 1050–1700 €/kWbiocoal for the smaller plants and 700–1100€/kWbiocoal for the larger plants. The specific investment for HTC-4.00 is the highest,due to the highest ballast of ash and moisture. Producing 1 MJ biocoal from MOWrequires 2.6 times the feedstock mass as does HTC from wood chips. This leads to largerequipment and consequently higher capital expenditure.

Biocoal production costs range from 4.9 €/GJ for HTC-4.00-s to 25.6 €/GJ for HTC-3.00-s-G for the smaller plants and from −0.4 to 14.3 €/GJ for the larger plants. Theproduction cost for HTC-4.00-m is negative because the remuneration for disposing of theMOW exceeds the cost of operating the HTC plant.

The biggest contributor to the total costs is the carrying charges (i.e. the costs relatedto the investment) except for HTC from wood, where the purchase of wood chips is themajor cost contributor. Labour cost are the second highest contributor for the smallerplants. For the larger plants processing MOW or PGW, the feedstock transport cost isthe second biggest cost factor.

120


Tabl

e4.

27:P

lant

capa

citi

es,s

peci

ficin

vest

men

tan

dle

veliz

edbi

ocoa

lpro

duct

ion

cost

sfo

rth

eH

TC

base

desi

gn.

HT

C-

1.00

-s2.

00-s

3.00

-s4.

00-s

5.00

-s3.

00-s

-G1.

00-m

2.00

-m3.

00-m

4.00

-m5.

00-m

*

feed

stoc

ken

ergy

[MW

HH

V]

10.9

9.5

9.5

7.9

11.9

9.5

54.3

47.5

47.5

39.4

26.6

bioc

oale

nerg

y[M

WH

HV

]8.

87.

47.

25.

79.

17.

243

.837

.136

.028

.320

.5

feed

stoc

km

ass

[kt F

M/a

]28

.028

.046

.746

.740

.046

.714

0.2

140.

223

3.6

233.

690

.0

bioc

oalm

ass

[kt F

M/a

]9.

18.

58.

28.

69.

18.

245

.742

.541

.043

.020

.5

spec

ific

TC

I[€

/kW

bio

coal

]11

7113

8914

0217

8710

5714

3775

288

586

511

0685

1

carr

ying

char

ges

[€/G

J]6.

547.

767.

829.

985.

898.

034.

204.

944.

836.

174.

75

labo

ur[€

/GJ]

4.00

5.46

5.62

6.80

0.26

4.65

1.39

1.97

2.03

2.40

0.15

elec

tric

ity

[€/G

J]1.

171.

371.

641.

931.

991.

641.

171.

371.

641.

931.

99

shel

ls&

fibre

s[€

/GJ]

0.09

0.09

O&

M,m

ater

ial

[€/G

J]1.

171.

421.

401.

821.

161.

400.

730.

890.

861.

120.

91

othe

rop

erat

ing

cost

[€/G

J]0.

020.

020.

020.

030.

020.

020.

020.

020.

020.

030.

02

biom

ass

1)[€

/GJ]

6.32

-3.0

6-5

.24

-16.

710.

539.

146.

32-3

.06

-5.2

4-1

6.71

0.53

biom

ass

tran

spor

t[€

/GJ]

0.27

0.83

1.45

1.08

0.72

0.47

3.08

4.24

4.64

bioc

oalt

rans

port

[€/G

J]1.

661.

66

bio

coal

cost

[€/G

J]19

.48

13.8

012

.72

4.92

11.6

025

.60

14.3

09.

218.

38-0

.41

10.1

11)

incl

udin

gco

stfo

rdr

ym

atte

rlo

ssdu

ring

stor

age

121


With the current assumptions on labour requirement, the smaller plants require 3 plantoperators and 1 unskilled worker per shift for the biomass handling, and the larger plantshave 4.5–5 operators and 2 workers. Plants for waste processing require an additional 1–2workers for sizing and sorting the feedstock. A high degree of automation may significantlyreduce the labour demand. The labour cost for the plant processing EFB are low due tothe lower wages in Malaysia.

The biocoal production costs are 4.3–5.3 €/GJ lower for the bigger plants due to economyof scale effects reducing both the capital investment and the labour requirement.

The waste water treatment is a significant cost factor due to the high electricity demandfor the aeration. Treatment costs comprising carrying charges for the treatment plant,maintenance costs and electricity amount to 5 €/m3 for HTC-3.00-s and HTC-3.00-m.This corresponds to 1.1–1.2 €/GJ, or 9–13% of the overall biofuel production cost. ForHTC-5.00 with EFB, the cost of waste water treatment rises to 1.7–1.8 €/GJ due tothe higher concentration of dissolved compounds. If the palm oil mill does not have aconnection to the electricity grid, it needs to be confirmed that the additional electricitydemand of the HTC process can be covered by the existing steam turbine at the palm oilmill.

In the simulation cases HTC-1.00, HTC-2.00, HTC-3.00, and HTC-4.00 the boiler forsteam production is fired with biocoal pellets. For small HTC plants in particular, it maybe more feasible to use a natural gas boiler instead due to its lower investment cost andeasier commissioning.

HTC from EFB offers an additional source of income if HTC treatment replaces thedumping of the EFB and if the avoided GHG emissions can be monetized. Figure 4.9 showsthe net biocoal production cost as a function of the CO2 price. Wood pellet productioncosts from WP-1.0-m-FR are shown for comparison. With the March 2012 CO2 price of 5€/t for Certified Emissions Reductions under the Clean Development Mechanism (CDM),factoring in the avoided emissions leads to a 7–9% reduction in biocoal price. With a CO2price of 20 €/t, the biocoal production cost are reduced by 30–34%. These calculationsonly take into account the avoided emissions during the production of the biocoal and notthe replacement of fossil coal with biocoal in a power station. A CO2 price of about 18€/t is required to make HTC-5.00-m* competitive with wood pelletizing.

4.5.11 Exergoeconomic analysis

An exergoeconomic analysis was performed for the simulation cases HTC-1.00-s and HTC-3.00 -s. These two cases are typical of the HTC from expensive biomass crops on the onehand and from waste biomass with a negative feedstock cost (remuneration for disposal)on the other hand. Figure 4.10 shows the results for the cost of exergy destruction CD

and the value Z, representing the carrying charges and operating and maintenance cost ofthe equipment, for the most important plant components. The values are normalized to1 GJ of biocoal output to make the two cases easier to compare. It is interesting that theCD and Z of most plant components are quite similar for both cases, although the PGWin HTC-3.00 has a negative cost while the wood chips in HTC-1.00 are expensive andconstitute the biggest cost contributor of the biocoal production. This is best explained bythe fact that only 16–19% of the chemical exergy of the biomass is actually converted in the

122


Figure 4.9: Production costs of biocoal from EFB, factoring in certificates for avoidedGHG emissions from EFB dumping. The production costs of wood pellets are shownfor comparison.

HTC reactor. The remainder passes through the reactor and all downstream components,and its cost is passed onto the final biocoal product C59, without much impact on theCD of the components. All the components responsible for pressurizing, preheating andrecovering thermal energy from the biocoal slurry use physical exergy or electricity as thefuel and are therefore only mildly affected by the cost of the chemical exergy in the flowstreams that they are processing.

The negative cost of the feedstock in HTC-3.00 leads to a negative cost for the wastewater before treatment (steam 94) and many flow stream which contain biocoal or dis-solved byproducts. These negative costs can be interpreted as a liability related to thesematerial streams, since the waste water needs treatment before it can be released to theenvironment. However, the waste water costs are more meaningful when expressed perunit of volume rather than per unit of exergy. The cost of stream 94 before treatmentcorresponds to −1.8 €/m3, while the treatment cost including carrying charges, mainten-ance cost and electricity consumption amounts to 5 €/m3. Thus, a positive value resultsfor the treated water.

The biggest difference between the two plants is the cost of exergy destruction in theboiler. The cost of the biocoal, which serves as a boiler fuel, is 2.7 times higher whenproduced from wood chips than when produced from PGW. It would actually be betterto burn raw wood chips in the boiler rather than carbonize them first. This however, isonly possible with high quality biomass such as wood chips, not with wet waste biomasssuch as PGW-70.

The components with the highest sum of CD + Z are the drier (K22), the HTC reactor(K14), the drier heat exchanger (K32), the filter press (K21), the pellet press (K46) andthe boiler (K25).

123


-1

0

1

2

3

4

5

6

7

8

9

10

11

K14

HTC

reac

tor

K22

drie

r

K32

drie

r HX

K21

filte

r pre

ss

K46

pelle

t pre

ss

K25

boile

r

K3 s

lurry

HX

K4 s

lurry

HX

K15

offg

as H

X 1

K26

FW p

rehe

ater

K36

FW p

rehe

ater

K41

offg

as H

X

K12

air p

rehe

ater

K5 s

team

/ sl

urry

mix

K7 s

team

/ sl

urry

mix

K1 s

lurry

mix

ing

tank

K9 s

team

/ sl

urry

mix

K13

stea

m /

slur

ry m

ix

K2 s

lurry

pum

p

K6 s

lurry

pum

p

K8 s

lurry

pum

p

K10

slur

ry p

ump

K17

flash

tank

K18

flash

tank

K19

flash

tank

K20

flash

tank

K38

drie

r fan

K24

cool

er

K33

cool

er

K39

mix

K29

mix

K42

cool

er

HTC-3.00-s

0

1

2

3

4

5

6

7

8

9

10

11

K14

HTC

reac

tor

K22

drie

r

K32

drie

r HX

K21

filte

r pre

ss

K46

pelle

t pre

ss

K25

boile

r

K3 s

lurry

HX

K4 s

lurry

HX

K15

offg

as H

X 1

K26

FW p

rehe

ater

K36

FW p

rehe

ater

K41

offg

as H

X

K12

air p

rehe

ater

K5 s

team

/ sl

urry

mix

K7 s

team

/ sl

urry

mix

K1 s

lurry

mix

ing

tank

K9 s

team

/ sl

urry

mix

K13

stea

m /

slur

ry m

ix

K2 s

lurry

pum

p

K6 s

lurry

pum

p

K8 s

lurry

pum

p

K10

slur

ry p

ump

K17

flash

tank

K18

flash

tank

K19

flash

tank

K20

flash

tank

K38

drie

r fan

K24

cool

er

K33

cool

er

K39

mix

K29

mix

K42

cool

er

C, D

,k+Z

k[€

/GJ P

,tot]

CD,k, HTC-3.00

Zk, HTC-3.00

CD,k, HTC-1.00

Zk, HTC-1.00

Figure 4.10: Cost of exergy destruction CD,k (bottom bar) and cost associated with theequipment items Zk (top bar) for HTC-3.00-s (left, green) and HTC-1.00-s (right, or-ange). Values normalized to 1 GJ of biocoal output.

The exergoeconomic factor f for the HTC reactor is 88–91%, indicating that investmentand operating costs are dominant rather than the cost of exergy destruction. The onlyway to reduce the investment cost is to decrease the volume of the reactor by decreasingthe residence time or by decreasing the water to biomass ratio of the slurry. Both optionsare explored in section 4.5.11.

The exergoeconomic factor for the drier, which has the highest CD + Z value of all plantcomponents, is 26–31%. The exergoeconomic factor of the drier heat exchanger for heatingup the drying air is 10–13%. Both components are dominated by the cost of exergydestruction. A design employing a more efficient drying technology, SSD, is analyzed insection 4.5.11. SSD has also been shown to reduce the exergy loss via the exhaust gas (seesection 4.1). In HTC-3.00-s, the cost of the drier exhaust gas is 20 €/h, equivalent to 6%of the total biocoal production cost.

The filter press and pellet press are both dominated by investment costs. There may bepotential for improvement by tailoring the equipment to the properties of the biocoal.However, that task falls outside scope of this work.

In HTC-3.00-s, the second biomass slurry preheater (K4) has a relatively high CD + Zvalue. Its exergoeconomic factor is 16% and its exergetic efficiency 62%. To decrease theexergy destruction, the mean temperature difference, which is now 52°C, would have tobe decreased. However, the biomass slurry must not be heated to a temperature above100°C in heat exchangers, in order to avoid fouling problems caused by the decomposingbiomass. The temperature of the inlet steam is set by the pressure of the flash stage.

124


Introducing more flash tanks would allow better matchings between hot and cold streams,but would also make the plant more complex.

Detailed results of the exergoeconomic analysis for HTC-3.00-s and HTC-1.00-s are givenin section B.4.6.

4.5.12 Parameter studies and alternative designs

The sensitivity of efficiency and cost to operating conditions and changes in the plantconfiguration is assessed, using HTC-3.00-s as the reference case. In a first step, keyprocess parameters including the degree of carbonization, the reaction temperature andpressure, and the water content of the biomass slurry and dewatered biocoal are varied. Ina second step, the influence of the flowsheet design on the thermodynamic and economicperformance of the plant is analyzed by changing the preheating scheme, dewatering anddrying of the biocoal. The simulation cases and key results are summarized in Table 4.28.The details are discussed in the following.

4.5.12.1 Parameter studies

Simulation cases HTC-3.01 and HTC-3.02 represent a weak and strong degree of car-bonization (see section 4.5.3). Cases HTC-3.03 and HTC-3.04 model the same degreeof carbonization as the base case, but achieve this with different pairings of temperatureand residence time. The efficiencies for the simulation cases with a reactor temperature of210°C, HTC-3.01 and HTC-3.04, are approximately the same as the base case. The effi-ciencies for the cases with a reactor temperature of 230°C are approximately 2 percentagepoints lower than that of the base case. As discussed in section 4.5.3, a higher reactiontemperature increases the amount of steam lost with the gaseous byproducts. In caseHTC-3.02, the stronger carbonization leads to a lower energy yield and to more gaseousbyproducts, and consequently a higher steam loss from the reactor. These effects resultin the lowest efficiency of the four cases at 70.2%.

The investment costs are strongly dependent on the reactor volume, which, in turn, isrelated to the residence time of the HTC reaction. The strongly carbonized biocoal is18% more expensive than the biocoal from the base case, and the weakly carbonizedbiocoal is 7% cheaper. Whether the better product quality achieved by a more severecarbonization justifies its higher cost requires further investigation, including combustionand gasification trials. The biocoal costs from HTC-3.03 and HTC-3.04 are 3% lowerand 6% higher compared to the base case, respectively. This indicates that, to someextend, shortening the residence time at the expense of a higher operating temperature isfavourable.

As discussed in section 4.5.3, the water vapour leaving the reactor with the gaseousbyproducts can represent a significant energy loss. This loss can be limited in two ways:by increasing the reactor pressure, and by constructing the reactor in a way that allowsdirect heat transfer between the biomass entering the reactor and the gaseous byproductsleaving the reactor. The former is modelled in case HTC-3.05 with a reactor pressure of 30bar. The latter is represented by HTC-3.09, where the gaseous byproducts are cooled to

125


Table 4.28: Energetic efficiency ηHHV, TCI, specific TCI per unit of biocoal and levelizedbiocoal production costs cbiocoal for different operating parameters and flowsheet designs.All cost data is for plants of small capacity (denoted -s). The basecase conditions aregiven in section 4.5.4.

ηHHV TCI specific TCI cbiocoal description[–] [M€] [€/kWbiocoal] [€/GJHHV]

parameter variationsHTC 3.00 72.6% 10.1 1402 12.72 base case (220°C, 4 h)HTC-3.01 72.2% 9.4 1316 12.17 210°C, 3 hHTC-3.02 70.2% 12.7 1824 15.50 230°C, 8 hHTC-3.03 70.9% 9.6 1367 12.65 230°C, 2 hHTC-3.04 72.7% 11.5 1597 13.86 210°C, 8.4 hHTC-3.05 73.5% 10.1 1384 12.58 preactor =30 barHTC-3.06 69.0% 10.5 1533 13.73 wfilter press =50%HTC-3.07 73.4% 9.8 1347 12.34 wfilter press =30%HTC-3.08-p1 73.9% 7.9 1075 10.31 wslurry =70%, pumpHTC-3.08-p2 73.0% 8.8 1209 11.89 wslurry =70%, plug screwHTC-3.09 73.8% 9.8 1337 12.22 Tgas=196°CHTC-3.10 73.9% 9.7 1324 12.14 Tgas=168°CHTC-3.11 73.9% 9.9 1346 12.32 Tgas=175°C, preactor =30 barHTC-3.12 73.6% Tambient=30°CHTC-3.13 64.6% Tambient=-10°Calternative designsHTC-3.20 75.5% 10.7 1413 13.09 atmospheric SSDHTC-3.30 77.7% 11.5 1453 13.68 atm. SSD, steam compressionHTC-3.40 70.8% pressurized SSDHTC-3.50 73.4% 10.7 1474 13.46 HT dewateringHTC-3.60 73.7% 9.9 1345 12.96 HT dewatering, plug screwHTC-3.61 73.2% inert gasHTC-3.62 73.5% wfilter press =30%HTC-3.70 79.3% 8.8 1104 10.51 heat exchangersHTC-3.80 46.0% 9.5 2075 19.20 no heat recoveryHTC-3.90 78.5% 10.7 1394 11.94 AD of waste water

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196°C by direct heat transfer in a counter-flow arrangement with the biomass slurry. CaseHTC-3.10 models an even stronger internal heat transfer where the gaseous byproductsare cooled to 168°C. Case HTC-3.11 combines internal heat transfer with an elevatedreactor pressure of 30 bar. The amount of steam in the gaseous byproducts is reducedby 68–96% by these measures. The higher reactor pressure in case HTC-3.05 leads toan efficiency increase by 0.9 percentage points and a decrease of biocoal production costby 4%. All three cases with internal heat transfer between the biomass slurry and thegaseous byproducts lead to an increase in efficiency of 1.2–1.4 percentage points. The TCIis 2–4% lower than in the base case, and the biocoal production costs drop by 6–7%. Thedecrease in TCI is mostly due to smaller heat exchanger surfaces for condensing vapourfrom the gaseous byproducts (K15, K26, K41). When combined with the internal heattransfer, the higher reactor pressure in HTC-3.11 does not lead to an increase in efficiencycompared to HTC-3.10. The TCI and biocoal production costs are slightly higher due tothe more expensive reactor and the slurry pump specifications adjusted to the higher op-erating pressure. The extent to which heat transfer between the gaseous byproducts andthe biomass entering the reactor can be realized in practice depends on the specific reactordesign and requires further investigation.

In simulation cases HTC-3.06 and HTC-3.07, the water content of the biocoal press cakeexiting the filter press is varied. A water content of 50% leads to a decrease in efficiencyby 3.6 percentage points. The TCI increases by 4% due to the higher capacity of the drierand boiler, and the biocoal production costs increase by 5%. A water content of 30%,which corresponds to the highest degree of dewatering achieved with laboratory presses,leads to a decrease in biocoal production costs by 6%.

The water content in the biomass slurry has an even stronger effect on the economicperformance, since a lower ballast of water flowing through the process makes most plantequipment smaller in size and therefore cheaper. In simulation case HTC-3.08-p1 with asolid matter content of 30% in the biomass slurry, the TCI is 22% lower and the biocoalproduction costs are 21% lower than in the base case. However, a higher dry mattercontent in the slurry makes the pressurization more demanding, and pumps may no longerbe applicable. Therefore, the economic calculation was performed additionally with datafor a special plug screw feeder which can handle solid matter contents of up to 70% (HTC-3.08-p2 ). Although the plug screw feeder is 550 k€ more expensive than the piston pumpin the base case, the biocoal production costs are 9% lower. It has to be pointed out thatthere is a high degree of uncertainty regarding the performance, cost and availability of thevarious pressurization systems. However, the analysis indicates that the amount of watercarried through the system has a big impact on the economic performance, and investingin pressurization systems that can deal with a high solid content is likely to pay off.

While the ambient temperature is generally 15°C in the simulation models, in the casesHTC-3.12 and HTC-3.12 it is set to 30°C and −10°C, respectively. The ambient tem-perature of 30°C leads to an increase in efficiency of 1.1 percentage points. The ambienttemperature of −10°C, and frozen biomass, lead to an efficiency decrease of 8.0 percentagepoints.17

17It is assumed that all the biomass moisture is frozen. However, chemically bound and capillary watermay freeze at much lower temperatures or not at all. For some lignites, up to 50% of the moisture doesnot freeze [339]. In this case, the efficiency loss would be less.

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The carbon yield for all the considered cases lies between 71–79%, being lowest for caseHTC-3.01 with strong carbonization. The carbon yield is mostly dependent on the amountof biocoal that needs to be burned in the boiler, which ranges from 7–19% of the grossbiocoal production.

4.5.12.2 Alternative flowsheet designs

In this section, seven different flowsheet designs for the HTC plant are analyzed, withchanges made to the biomass pressurization and preheating scheme for the biomass slurryand the biocoal dewatering and drying. The base case HTC-3.00 employs a complex pre-heating scheme where steam is mixed with the biomass at several pressure levels, thusintertwining pressurization and preheating of the biomass. This may cause operabilityissues, and further complicate the pressurization of the biomass, which is technically de-manding in any case. A design where the biomass can be preheated in one stage wouldbe more attractive for operability, but requires a different solution for the preheating.

As shown in section 4.5.12.1, the efficiency of the mechanical dewatering is of great signi-ficance to the overall efficiency and cost of the biocoal production. In the base case, thebiocoal is dewatered after the de-pressurization of the coal slurry, thus at a temperaturebelow 100°C. Dewatering at the outlet of the reactor, at more than 200°C, may facilitatea higher degree of dewatering. However, it is technically more complicated, because itmust take place in a pressurized atmosphere. Designs with high temperature dewateringare analyzed in simulation cases HTC-3.50 and HTC-3.60.

The exergy analysis and exergoeconomic analysis of the base case indicate that the drierhas a major impact on the exergy destruction and economic performance of the system.Therefore, several flowsheet designs employing superheated steam drying (SSD) are in-vestigated.

Table 4.29 gives an overview of the analyzed flowsheet designs. The results are summarizedin Table 4.28. The respective flowsheets, if not shown in this section, can be found insection B.4.7.

Table 4.29: Alternative designs for HTC plants.

pressurizing heat recovery dewatering dryingHTC-3.00 4 stages, 15% DM flash/mix LT LT waste heatHTC-3.20 4 stages, 15% DM flash/mix LT atm. SSD, SCHTC-3.30 4 stages, 15% DM flash/mix, SC LT atm. SSD, SCHTC-3.40 4 stages, 15% DM flash/mix LT press. SSDHTC-3.50 4 stages, 15% DM flash/mix HT LT waste heatHTC-3.60 1 stage, 30% DM hot filtrate rec. HT LT waste heatHTC-3.70 1 stage, 15% DM HX LT LT waste heatHTC-3.80 1 stage, 15% DM none LT LT waste heat

DM: dry matter content, HT: high temperature, LT: low temperature, SC: steam compression, press.:pressurized, atm.: atmospheric, rec.: recirculation, HX: heat exchanger

In HTC-3.20, SSD at atmospheric pressure is employed. The exhaust steam from the drieris compressed to 6 bar and utilized for preheating the biomass slurry and as heating steam

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for the drier. The efficiency is 2.9 percentage points higher than that of the base case.However, the biocoal production cost are also 3% higher due to higher investment costs.The amount of biocoal required for steam generation is reduced by 56%. However, becauseof the steam compressors, the electricity consumption increases to 72 kJ/MJbiocoal, 25%higher than in the base case. There is a surplus of steam at 1 bar which cannot be utilizedand is discharged to the condenser.

HTC-3.30 employs SSD drying as in HTC-3.20, but the surplus 1 bar steam is compressedto 31 bar and mixed with the biomass slurry. No biocoal boiler is required in steady-stateoperation, but a gas boiler for start-up is included in the investment cost estimate. Aregenerative incinerator is employed to oxidize the gaseous byproducts of the reactor.Compared to HTC-3.20, this design increases the efficiency by another 2.2 percentagepoints, and the biocoal production cost by another 4%. The electricity consumptionamounts to 87 kJel/MJbiocoal.

In HTC-3.20 and HTC-3.30, the steam compressors contribute 418 k€ and 787 k€, re-spectively, increasing the TCI by 6% and 14% compared to the base case. Moreover, thesteam from the drier may contain condensable organic compounds which may impede thesteam compression in practice. HTC-3.40 therefore represents a design with SSD wherebysteam compression is avoided. The SSD is operated at elevated pressure. The efficiencyis 1.8 percentage points lower than in the base case, because the 15 bar heating steamconsumed by the drier reduces the amount of steam available for preheating the biomassslurry, which has to be compensated for by additional biocoal combustion. Because of thedrier exhaust steam, there is more steam at 4 bar than can be used for preheating. Thissurplus steam is therefore discharged to the condenser.

The flowsheets of HTC-3.20, HTC-3.30 and HTC-3.40 are given in Figure B.2, Figure B.3and Figure B.4.

In HTC-3.50, the biocoal is mechanically dewatered at the reactor outlet at high pressureand temperature.18 The subsequent de-pressurization of the biocoal leads to a furtherreduction in the water content to 30% by flash evaporation. The flowsheet is shown inFigure B.5. The efficiency is 0.8 percentage points higher than in the base case. Thebiocoal production cost are 6% higher than in the base case due to the higher investmentcost of the dewatering press operating at high pressure. There is a high uncertaintyabout the technical feasibility and the investment costs of the dewatering at 25 bar and220°C. Based on the assumptions made here, the press is twice as expensive as that inthe base case due to the larger mass flow being processed and its high operating pressureand temperature. The additional investment for the dewatering press is partly offset bysavings in downstream components, primarily the drier.

HTC-3.60 combines high temperature dewatering with single-stage biomass pressuriza-tion. The flowsheet is shown in Figure 4.11. The biomass is pressurized with a plug-forming feeder without adding water. After pressurization, it is mixed with water re-covered from the dewatering press to obtain a solid matter content of 15% in the reactor.The amount of water needed to cover the biomass with water in the reactor may be lower18The same dewatering performance is assumed as in the base case, leading to a water content of 40% for

the dewatered biocoal. For mechanical-thermal dewatering of lignite, the lower density and viscosity ofwater at higher temperatures lead to an improved dewatering behaviour [202, page 16]. Whether thisis also the case for HTC biocoal needs further investigation.

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biocoal

condensate

reactor

pelletpress

flue gas

filter press

air

6

54

31

20

20

19

48

74

53

30

58 66

10

7

5 67

80 86

56

62

73

22

23

61

28

63

63

8564

70

59

70

100

60

29 75

90

8

8

95

95

9

32

34

46

55

44

36

68+99

42

K2 K3

K28

K24

K45

K16

K29

K21

K37 K44

K46K22

K18

K40

K27

K12

K39

K38K32

K26

K25

K42

K43

K13

K14

K17

21 4

air

drier

A

A

B

B

C

C

D

E

E

D

biomass

27

1

K33

W1

W9

W7

W4

W8

W3

wastewater

aeration

discharge

94

W14

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

ash72

Figure 4.11: Flowsheet of HTC-3.60.

in practice, depending on the bulk density of the biomass. The reactor is operated at 30bar. The efficiency is 1.2 percentage points higher than in the base case. The biocoal pro-duction costs are 2% higher, due to the higher electricity consumption of the plug-formingfeeder compared to a pump, and the higher operating and maintenance costs for high wearcomponents.

Two parameter studies are conducted for this flowsheet design. HTC-3.61 investigatesthe influence of air entering the system with the biomass. 14 g air per kg biomass isassumed based on data for a piston feeder for wood chips [251]. The additional gas in thereactor increases the steam in the gaseous phase by 33%, leading to an efficiency loss of0.5 percentage points. Other systems for feeding biomass against pressure can result inmuch higher amounts of gas entering the reactor. For example, Drift reports that for alock hopper system, the amount of inert gas is 20 times as high as for the piston feeder[251]. Because of the negative effect on the plant efficiency, feeding systems that result ina high amount of air or other gas entering the reactor should be avoided.

HTC-3.62 represents the same design as HTC-3.60 with a higher degree of mechanicaldewatering. Other than in the base design, the dewatering to 30% water content doesnot lead to an increase in efficiency.19 There is a surplus of steam at 1 bar anyway,and the higher degree of dewatering simply leads to more steam being discharged to thecondenser. However, more efficient dewatering will decrease the required drier capacityand may therefore lead to an economic benefit.

The base design does not employ heat exchangers at temperatures exceeding 100°C dueto the fouling problems reported for the hydrothermal treatment of peat [200]. However,the use of indirect heat transfer at temperatures up to 207°C with a specially developedheat exchanger was employed in the DeLaval’s Process, another peat upgrading scheme

19The slight decrease in efficiency by 0.1 percentage point is due to the higher degree of dewatering, whichleads to more dissolved organic compounds remaining in the water rather than the press cake.

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condensate

reactor

filterpress

naturalgas

flue gas

air

666

17

31

54

53

4

24

25

8

10

6780 88 89

62

18

28

28

85

85

23

23

30

30

20

2022

22

64

35

60

95

32

34

44 36

68+99

39

39

42

K2 K3

K4

K28 K24K16

K21

K37

K22

K20

K27

K38

K42

K14

air

drier

A

A

B

B

C

C

biomass

2758

1 K1

K19

K18

K17

K33

3

2W1

W2

W7

W8

incinerator

K25

D

D

wastewater

aeration

discharge

94

W14

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

biocoalpelletpress

5929K46

W9

K44

Figure 4.12: Flowsheet of HTC-3.70.

[154]. A design which employs heat exchangers (K3, K4) to preheat the biomass slurryto the reactor inlet temperature of 191°C is therefore analyzed in case HTC-3.70. Theflowsheet is shown in Figure 4.12. Water at 40 bar is used as a thermal fluid to transferthermal energy from the biocoal slurry to the biomass slurry. Gaseous byproducts fromthe reaction (39, 67) are combusted in an incinerator. The thermal energy released is usedfor heating the drying air. Additional thermal energy for the drying is recovered fromthe waste water cooling (K 19), the depressurized biocoal slurry (K18) and the reactoroffgas (K17). A small amount of natural gas amounting to 0.01 MJ/MJbiocoal is utilizedto support the combustion of the low calorific gaseous byproducts and provide sufficientthermal energy for drying the biocoal. At 79.3%, the energetic efficiency is the highest ofall analyzed plant designs, 6.7 percentage points higher than that of the base case. Theinvestment is 13% lower than that of the base case due to the simpler plant design andthe biocoal production costs are 17% lower.

In order to quantify the importance of the heat recovery, case HTC-3.80, without heattransfer from the biocoal to the biomass slurry, is investigated. The flowsheet is shown inFigure B.6. The biocoal slurry is de-pressurized to 1 bar in one step. Some steam is utilizedto provide thermal energy to the drier, while the remainder is discharged to the condenser.The efficiency is only 46%, because more than 40% of the produced biocoal is burned forsteam generation. The TCI is only 6% lower than for the base case. The lower costs forflash tanks, heat exchangers and and slurry pumps are almost compensated by higher costsfor the reactor,20 larger capacities for waste water treatment and the additional capacityof coolers to discharge the waste heat. The biocoal production cost is 51% higher than inthe base case. This demonstrates that HTC without efficient heat recovery is clearly notsensible.

Case HTC-3.90 with anaerobic digestion of the waste water is discussed in section 4.5.13.20The reactor volume increases as a result of the higher steam consumption.

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In conclusion, the two cases which cover the entire steam demand with internal heatrecovery, HTC-3.70 and HTC-3.30, have the highest efficiencies of 78–79%. They also havethe maximum carbon recovery rate of 85%. Designs employing heat exchangers to preheatthe biomass slurry and cool the biocoal slurry are potentially amongst the most efficientand cheapest. Therefore the potential to use heat exchangers should be explored further.This requires comprehensive trials to investigate the fouling and clogging characteristicsof the biomass and biocoal slurries at different temperatures and their effects on differenttypes of heat exchangers.

Excessively long residence times are detrimental, since the increased reactor volume markedlyincreases the investment costs. For this reason, “over-carbonization”, meaning carbonizingbeyond that required to reach the desired properties of the biocoal, should be avoided.

The evaporation of water in the reactor should be minimized by employing reactor designsthat cool the offgas in direct heat transfer with the incoming biomass, by operating thereactor at a high pressure, and by preventing air or other gases from entering the reactor.The biomass feeding and pressurization system, besides from being able to cope with ahigh solids content, should minimize the amount of inert gas or air entering the reactor.This may further constrain the choice of suitable pressurization systems.

A high degree of mechanical dewatering decreases the investment cost of the drier. Whetherit also improves the efficiency of the plant depends on whether the saved 1 bar steam canbe utilized for preheating the biomass slurry.

The efficiency for most of the designs and operating parameters analyzed lies within therange of 70–78%, with biocoal production costs of 10–15 €/GJ. A high degree of heatintegration between the biocoal and biomass slurry is essential.

The overall auxiliary energy demand, namely electricity and boiler fuel, ranges between6–20% of the gross biocoal energy for different feedstocks and plant designs. This in ingood agreement with published information from developers of HTC plants, who reportvalues between 6–19% [174, 195, 340].

More details on the analyzed cases are provided in section B.4.7, including a breakdown ofthe investment costs and biocoal production costs for all the analyzed cases, and flowstreamdata and detailed equipment lists for the cases HTC-3.30 and HTC-3.60.

4.5.13 HTC with anaerobic digestion of the wastewater

The waste water from the HTC coal dewatering carries a high load of organic and mineralcompounds. The dissolved organic compounds represent a major energy loss from theHTC plant, as discussed in section 4.5.5. Moreover, the waste water requires remedialtreatment. In contrast to aerobic treatment, anaerobic digestion can recover part of thisenergy in the form of methane. The resulting biogas can be combusted in the boilertogether with the reactor offgas and biocoal.

The HTC waste water is rich in short-chain organic acids, including formic and aceticacid. These substances are intermediate products of anaerobic digestion, which can bequickly converted to methane, without the slower preceding step of hydrolysis. Wirthconducted AD experiments with HTC waste water from maize ensilage21 and achieved21The HTC was conducted at 220°C for 6 hours.

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methane rates of up to 650 l/kgTOC and a COD removal rate of 50–80% [220, 338]. Theoperating temperature was 37°C and the OLR was 1.0 kgCOD/m3/d.

Besides the easily degradable short-chain organic acids, more persistent substances likephenol and 5-hydroxymethylfurfural have also been detected in the HTC process water[162]. Wirth did not observe any obvious inhibiting effects on the microorganisms duringhis six week experiments, but depending on the composition of the process water, partialinhibition may occur [220]. Since these persistent substances are not degraded during AD,a subsequent aerobic treatment step would seem necessary. Another option is pretreatmentof the waste water upstream of the AD step to diminish phenols and make the waste watermore easily degradable [220].

In this section, the integration of anaerobic digestion of the waste water into the HTCbase case HTC-3.00 -s is analyzed. The case with AD is referred to as HTC-3.90-s.

4.5.13.1 Simulation model for the anaerobic digestion of the waste water

The HTC plant produces 6321 kg/h of waste water, of which 64% is press water from themechanical dewatering and the remainder is condensate from the flash tanks and fromthe reactor offgas. The solid matter content is 4.1% for the press water and 0.4% for thecondensate, resulting in 2.8% for the mixture of both. The composition of the flow streamsis given in section B.4.11.1.

Based on [220], it is assumed that 80% of the acetic and formic acid is converted accordingto Equations 4.2 and 4.3.

CH3COOH → CH4 + CO2 (4.2)

4 HCOOH → CH4 + 3 CO2 + 2 H2O (4.3)

The degradation rate of the remaining organic compounds TOMres based on the experi-mental data [220] corresponds to 50% of the theoretical methane production and 45% ofthe theoretical CO2 production, according to Equation 2.4. These same ratios of actualto theoretical conversion are assumed for the simulation, resulting in a degradation rateof the TOMres of 46%. The results are summarized in Table 4.30.

The chemical oxygen demand (COD) is used in environmental chemistry to characterizewaste waters and the treatment efficiency regarding organic compounds. The COD denotesthe oxygen demand for a complete oxidation of all organic compounds contained in asample. It is usually determined experimentally by standardized procedures. For thewaste water in the simulation, the COD is estimated as the stoichiometric oxygen demandfor the complete oxidation of the organic compounds based on their elemental composition.

The HTC waste water from the simulation contains almost twice as much acetic acidthan that used in the experiments, but 37% less TOMres. Since acetic acid is more eas-ily degradable, this results in a 5% higher methane yield and a higher degree of TOCdegradation.

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Table 4.30: Anaerobic digestion of HTC waste water: simulation model and measureddata [220].

simulation [220]feedacetic acid [g/l] 10.49 5.26formic acid [g/l] 0.75 0.81TOC [g/l] 11 15COD [g/l] 35 41conversionTOC degradation 61% 50%COD removal rate 59% 50–80%CH4 production [lSTP/kgTOC] 683 650CH4 production [lSTP/kgODM] 365biogas production [lSTP/kgODM] 611methane volume fraction [–] 60% 65%digestateTOC [g/l] 4.22COD [g/l] 14.44

It should be pointed out that there is a high degree of uncertainty about the degradationof the remaining organic compounds TOMres. Their composition, and consequentiallytheir degradation behaviour, may strongly depend on the choice of HTC feedstock andoperating conditions.

For a hydrothermal peat upgrading plant employing a similar heat recovery scheme withseveral flash stages, it is reported that the phenol content of the condensate is higher thanthat of the press water [200]. Depending on the composition of both waste water streams,it may be better to subject only the press water to anaerobic digestion.

Another aspect that needs consideration is the low pH of the HTC waste water. Althoughacetic acid is easily degradable, a too low pH is harmful for methanogenesis. Wirth reportsthat the acidic nature of the HTC waste water posed problems in batch experiments, butnot during continuous operation [220]. However, some laboratory experiments reveal aceticacid concentrations of up to 34 g/l in the HTC process water [23], considerably higherthan that encountered by Wirth. It may therefore be necessary to dilute the waste wateror co-digest it with other substrates.

Wirth conducted the experiments with an OLR of 1 gCOD/l/d but concludes that it canlikely be increased under optimized conditions. Sensitivity analysis showed that economicfeasibility is reached at an OLR of 1.5 gCOD/l/d or greater [220]. For the anaerobicdigestion of raw biomass, OLR in the range of 2–3 gODM/l/d [102, page 40] are commonlyapplied.

In the simulation, an OLR of 1.5 gODM/l/d is assumed, corresponding to 2.6 gCOD/l/d.According to Equation 2.5, this results in a hydraulic retention time (HRT) of 13.2 days.

Simulation data related to the substrate and products of the anaerobic digestion arepresented in section B.4.11.1. Details on the digester design and heat loss are given insection B.3.4.

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4.5.13.2 Energy and carbon balance

Based on the HRT, the digester can be sized and the heat loss calculated. Three digesterswith a volume of 835 m3 each are required. The heat loss of 13 kW is compensated bythe thermal energy of the substrate, which enters the digester of a temperature of 41°C.Details on the heat loss calculations can be found in section B.3.4. No heating is requiredto maintain the digester temperature of 38°C at an ambient temperature of 15°C. Higherheat losses in winter can be overcome by taking the waste water from the HTC plant ata higher temperature. However, this may compete with other low temperature heat usessuch as the preheating of the combustion air.

The electricity consumption for the stirrers is assumed to be 3% [28] of the biogas energyand amounts to 15 kW.

The energy content of the produced biogas is 505 kW (HHV), reducing the energy lossfrom the dissolved compounds by 53%. The biogas is used to replace biocoal in the boilerand thereby reduced the amount of biocoal combusted by 68%. Moreover, the electricityconsumption of the aerobic waste water treatment is reduced by almost 60% from 243to 99 kW, because more than half of the dissolved organic matter is converted in theupstream anaerobic digestion.

The overall efficiency (HHV) is 78.5%, 6 percentage points higher than in the base caseHTC-3.00. The carbon yield is also high at 82.6%.

HTC-3.90 with anaerobic digestion has a similar efficiency to the cases without an auxil-iary boiler, namely HTC-3.30 with SSD and recompression of the flash steam and HTC-3.80 with heat recovery by heat exchangers. There is no use in combining AD withHTC-3.30 or HTC-3.80, since there is no need for combustion fuel and the AD gas cannotbe utilized on-site.

It is assumed that no gas treatment is required for the combustion of the biogas in theboiler. Based on the simulation, the biogas contains 1674 mg/m3 H2S, but as discussedin section 4.5.8, the H2S production in the HTC reactor may have been overestimated.

4.5.13.3 Economic performance

The digester, consisting of three tanks of 834 m3 each, adds to the investment costs by 0.51million €. It is assumed that the investment cost of the aerobic waste water treatmentplant is the same as for HTC-3.00-s, although in practice it may be lower due to thelower organic loading which needs to be degraded in this treatment stage. The investmentfor the overall waste water treatment, comprising anaerobic and aerobic stages, is 70%higher than in HTC-3.00-s. However, this additional investment is overcompensated bythe savings in electricity, making the total treatment costs 4% lower for HTC-3.90-s. Afurther advantage is the replacement of part of the biocoal fuelling the boiler by biogas,leading to biocoal production costs of 11.59 €/GJ, 6% lower than for HTC-3.00-s.

Details of the cost estimates are presented in section B.4.11.2.

There is a high degree of uncertainty about the cost of waste water treatment. Neverthe-less, the anaerobic digestion of the waste water seems an interesting option to recover some

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of the energy in the dissolved compounds, leading to a higher energy efficiency and carbonyield, and at the same time potentially lowering the biocoal production costs. When theenergy contained in the dissolved organics can be utilized, minimizing the dissolved com-pounds is no longer as crucial. Indeed, producing sufficient biogas to cover the entire heatdemand of the HTC plant without having to burn biocoal may offer advantages, since theboiler will be cheaper.

4.5.14 HTC integrated with CHP

As previously discussed, efficient heat recovery within the HTC plant is essential for bothits energetic efficiency and its biocoal production costs. Instead of utilizing the thermalenergy released during the cooling and de-pressurization of the products within the HTCplant, an interesting option would be to couple the HTC process with a rankine cycle CHPplant. In this case, steam and water at different temperature and pressure levels can beexchanged between the two processes and may thereby be utilized more efficiently. Thiswould also omit the need to combust part of the produced biocoal for steam production,since steam for the HTC reactor can be taken from a turbine extraction. This seems ap-pealing since the biocoal boiler in the standalone HTC plant has a high exergy destructionand low exergetic efficiency.

Integration of an HTC process with a simple rankine cycle CHP plant is investigated inthis section. The CHP plant is fired with raw wood and produces 20 MW of district heatat 90°C. The return temperature of the district heating water is 55°C. The steam turbineinlet conditions are 500°C and 80 bar. The HTC process has a capacity of converting 6.7t/h of PGW-70.

Three designs with combined heat, power and biocoal pellet production are investigated:CHPB-3.1 with a low degree of integration between the two processes, CHPB-3.2 with ahigh degree of integration, and CHPB-3.3 with a high degree of integration and SSD. Astandalone CHP plant plus standalone HTC-3.00 is used as the reference case to assessthe effects of integration on the thermodynamic and economic performance.

In CHPB-3.1, the HTC plant and CHP plant are basically separate plants. The design ofthe HTC process is mostly identical to the standalone HTC plant, but instead of a boiler,a steam turbine extraction at a pressure of 34.7 bar provides thermal energy to the HTCreactor.

The rationale behind CHPB-3.2 is the simplification of the heat recovery scheme forthe HTC process. The intertwined preheating and pressurization of the biomass slurryin HTC-3.00 may lead to operability issues and high costs. Stronger integration of theHTC process with the CHP plant can eliminate the need for this complex preheatingscheme. The flowsheet of CHPB-3.2 is shown in Figure 4.13. Flow stream data is givenin Table B.35. The biomass slurry is pressurized in one stage, preheated to 100°C by heatexchangers, and then heated to the the reactor inlet temperature by mixing with steam at29 bar (7). The thermal energy for producing this steam is provided by a steam turbineextraction (98). The biocoal slurry is depressurized in two steps, and steam is recovered at5 bar and 1 bar. Some of the recovered steam is used to preheat the biomass slurry (K3,K4), for feedwater preheating (K26, K36), and to supply the biocoal dryer. The remainderof the steam plus the thermal energy recovered from the gaseous byproducts is used for

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condensate

reactor

flue gas

filter pressair

76

4750

14

11

126

96

54

53

4

10

7

77

85

83

83

100

5

5

80 87

56

62

92

41

97

61

28

63

64

60

29

63

95

95

93

93

9

46

91

55

8

57

44

36

5868+99

30

42

K2 K3 K4

K48

K58

K28

K24

K45

K49K50

K52K53

K61

K60

K55

K54K11

K16K19

K31

K21 K37

K22

K20

K36

K27

K26

K38K32

K41

K56

K42

K13

K14

17 1871

drier

A

A

B

D

D

B

biomass

27

1 K1

K33

3

2W1

W7

W8

W3

wood

steamturbine

hotwater

heatingcondenser

heatingcondenser

ash 72

73

98

78

48

15 45

67 51

22

21

23

74

79W14+W15

W16

G

C

C

air

52

16

16

13

13

81

81

82

82

E

E

F

F

G

G

66

5934

32

35

K51

exhaustgas

wastewater

aeration

discharge

94

W18

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

biocoalpelletpress

75K46

W9

K44

Figure 4.13: Flowsheet of the combined HTC and CHP plant CHPB-3.2.

preheating the combustion air (K58, K41) and for district heat production (K57). TheHTC plant receives 2.3 MW of thermal energy from the CHP plant and returns 1.6 MWat a lower temperature level, still useful for district heat production and air preheating. Inthis way, the thermal energy recovered from the reaction products can be utilized withoutthe complex biomass slurry preheating scheme.

Simulation case CHPB-3.3 has the same configuration as CHPB-3.2, but employs a su-perheated steam drier, similar to HTC-3.20 and HTC-3.30. The exhaust steam of thedrier is used for district heat production (K57), therefore recompression is not necessary.This is an advantage compared to the standalone HTC plants with SSD, because steamcompressors are expensive and there may be operational issues due to impurities in therecovered steam.

The flowsheets for CHPB-3.1 and CHPB-3.3 are given in section B.4.12.

4.5.14.1 Energy and carbon balance

The energetic fuels and products and energetic efficiency of the analyzed systems are shownin Table 4.31. PGW accounts for 20% of the fuel, while the remainder is wood chips.

The standalone CHP plant has an electrical efficiency of 15.7% and a thermal efficiency of51.7%, resulting in an overall energetic efficiency (HHV) of 67.3%. The LHV-based overallenergetic efficiency is 83.8%.

The biocoal pellet production is 10% higher in the integrated designs than in the standaloneHTC configuration, because the combustion of part of the produced biocoal for steamgeneration is avoided. This is mostly compensated for by an increase in wood consumptionin the CHP boiler in CHPB-3.1 and CHPB-3.2. In CHPB-3.3, the wood consumption isslightly lower than in the standalone configuration.

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The net electricity production is 1% higher in CHPB-3.1 than in the standalone con-figuration. For CHPB-3.2, it is 2% lower. The omission of the complex biomass slurrypreheating scheme leads to an increase in steam demand of the HTC reactor by 300%compared with the standalone HTC plant. This steam is supplied by the steam extractionat 34 bar. After raising steam for the HTC process, the hot condensate is used for districtheat production. This steam flow therefore “circumvents” the low pressure turbine stage,in which most of the electricity is produced. In CHPB-3.3, the electricity production is 5%lower than in the standalone configuration. The steam mass flow through the low pressureturbine stage is determined by the district heat demand. Since the more efficient dryingtechnology frees up low pressure steam from the HTC process for district heat production,the steam mass flow through the turbine, and consequently its electricity production, arereduced.

The overall energetic efficiency is exactly the same for CHPB-3.2 as for the standaloneconfiguration, and 0.3 percentage points higher for CHPB-3.1. For CHPB-3.3, it is 1.2percentage points higher. The integration with the CHP plant therefore allows the complexheat recovery scheme in the HTC process to be omitted without sacrificing efficiency.

Only 15.5–17.2% of the feedstock carbon (wood plus PGW) ends up in the biocoal. Fora BECCS scenario, the CHP process would need to be equipped with carbon capture inorder to achieve an acceptable capture rate.

Table 4.31: Energetic fuels and products, important energy flows and energetic efficienciesof the integrated HTC and CHP systems.

standalone CHPB-3.1 CHPB-3.2 CHPB-3.3

fuel

wood [MW] 38.70 39.62 39.58 38.49

PGW-70 [MW] 9.51 9.51 9.51 9.51

products

electricity (net) [MW] 5.65 5.69 5.54 5.36

district heat [MW] 20.00 20.00 20.00 20.00

biocoal pellets [MW] 7.20 7.93 7.90 7.91

energy flows between CHP and HTC process

thermal energy from CHP to HTC [MW] – 0.66 2.35 2.07

district heat supplied by HTC [MW] – – – 0.67

air/FW preheating supplied by HTC [MW] – – 1.60 1.19

electricity production (>0) and consumption (<0)

high pressure turbine (K49) [MW] 1.68 1.72 1.77 1.72

low pressure turbine (K50) [MW] 4.51 4.51 4.33 4.18

feedwater pump [MW] -0.13 -0.13 -0.13 -0.13

HTC process [MW] -0.41 -0.41 -0.42 -0.40

energetic efficiency (HHV) [–] 68.1% 68.4% 68.1% 69.3%

carbon yield in biocoal [–] 15.5% 16.8% 16.8% 17.2%

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4.5.14.2 Exergy analysis

The results of the exergy analysis of the combined HTC and CHP plants are summarizedin Table 4.32.

The overall exergetic efficiency of the three integrated designs is 0.7–1.1 percentage pointshigher compared to the standalone configuration. The exergy destruction by steam/slurrymixing increases by 59% in the plant designs without the complex preheating schemeCHPB-3.2 and CHPB-3.3, because of the larger temperature difference between slurryand steam. Flashing the biocoal slurry in two stages instead of four increases the exergydestruction from flashing by 64%. The avoided exergy destruction by combusting part ofthe biocoal is mostly compensated for by an increase in exergy destruction in the CHPboiler.

Table 4.32: Exergy analysis results for the integrated HTC and CHP systems.


fuel

wood (DM) [MW] 40.65 41.56 41.53 40.38

PGW-70 (DM) [MW] 10.06 10.06 10.06 10.06

air + fuel moisture [MW] 0.27 0.28 0.28 0.24

products

electricity (net) [MW] 5.65 5.69 5.54 5.36

district heat [MW] 3.43 3.43 3.43 3.43

biocoal pellets [MW] 7.49 8.24 8.22 8.22

HTC process

exergy losses [MW] 1.11 1.06 1.10 1.05

drier exhaust gas [MW] 0.16 0.14 0.12 0.03

waste water [MW] 0.88 0.86 0.94 0.96

other [MW] 0.07 0.07 0.04 0.06

exergy destruction [MW] 1.98 1.46 1.68 1.59

HTC reactor [MW] 0.55 0.55 0.55 0.55

drier [MW] 0.17 0.15 0.15 0.06

boiler [MW] 0.48 – – –heat exchangers [MW] 0.10 0.09 0.15 0.14

steam-slurry mixing [MW] 0.12 0.12 0.20 0.20

flash tanks [MW] 0.08 0.09 0.13 0.13

other [MW] 0.47 0.46 0.50 0.51

CHP process

exergy losses [MW] 3.42 3.49 3.49 3.39

boiler exhaust gas [MW] 2.39 2.46 2.47 2.40

ash [MW] 1.03 1.02 1.02 0.99

exergy destruction [MW] 27.92 28.53 28.41 27.64

boiler, air preheater [MW] 25.24 25.82 25.66 24.95

steam turbine, pump [MW] 1.24 1.25 1.22 1.19

heating condensers [MW] 1.45 1.46 1.53 1.50

overall exergetic efficiency [–] 32.5% 33.5% 33.1% 33.6%

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4.5.14.3 Economic performance

The CBM for the HTC process are 0.6 M€ lower in the strongly integrated designs CHPB-3.2 and CHPB-3.3 compared to the standalone configuration, resulting in slightly lowerTCI for the overall system. CHPB-3.3 has the lowest investment cost at 25.4 M€. Detailson the investment cost estimates for the CHP process can be found in Table B.37. Theinvestment costs are summarized in Table B.36.

The total annual levelized revenue requirement (TRR) for the CHPB plants equals theproduction costs of the three products, namely electricity, district heat and biocoal pellets.Assuming that specific revenues for district heat and electricity are the same for all plantdesigns, the specific cost of the biocoal pellets (cbc) can be calculated using Equation 4.4,where W is the net annual electricity production, cw the renumeration for electricity feed-in, Qdh the annual district heat production, cdh the renumeration for district heat, andQbc the annual biocoal production

cbc =TRRCHP B − cwW − cdhQdh,CHP B

Qbc(4.4)

The remuneration for district heat cdh is calculated based on the required revenues for thedistrict heat from the standalone CHP plant and amounts to 33.7 €/MWh.

The annual cost flows of the standalone CHP plant and the CHPB plants are given inTable B.37.

Table 4.33 shows a breakdown of the biocoal production costs. The costs related to thebiocoal production for the integrated systems are calculated by difference with the stan-dalone CHP plant as indicated by Equation 4.5. Ci denotes the various cost items i suchas carrying charges, labour costs etc., and the subscripts CHP and CHPB denote thestandalone CHP plant and the combined heat, power and biocoal plant, respectively. Theelectricity cost in the standalone case is the purchased electricity for the HTC plant, whilein the integrated plants it is the lost feed-in revenues compared to the standalone CHPplant.

ci,bc =Ci,CHP B − Ci,CHP

Qbc(4.5)

The biocoal production costs from the integrated systems are 25–26% lower than those ofthe standalone HTC plant HTC-3.00. The cost savings are due to lower carrying charges(1.4–1.8 €/GJ) and labour costs (2 €/GJ).

In contrary to the standalone HTC plant design analyzed in section 4.5.12.2, employingSSD drying in the integrated plants leads to the lowest biocoal production cost, becausethe drier exhaust steam can be used for district heat production and expensive steamrecompression is avoided.

All in all, the integration of HTC with a CHP plant seems attractive, with lower cost, thesame efficiency and probably advantages regarding operability. However, locations wherethe availability of an adequate amount of PGW or similar biomass waste coincide withthe required district heat demand will be limited.

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Table 4.33: Investment costs of the CHPB systems, and levelized costs ci,bc assigned tothe biocoal production.

standalone CHPB-3.1 CHPB-3.2 CHPB-3.3TCI (HTC + CHP) [M€] 26.76 25.82 25.23 25.44carrying charges [€/GJ] 7.76 6.39 5.99 6.13labour [€/GJ] 5.62 3.58 3.59 3.59electricity [€/GJ] 1.64 1.05 1.46 1.96O&M, material [€/GJ] 1.40 1.27 1.18 1.20other operating cost [€/GJ] 0.02 0.03 0.03 0.02PGW incl. transport [€/GJ] -3.79 -3.44 -3.45 -3.45wood incl. transport [€/GJ] 0.00 0.62 0.60 -0.14total biocoal cost [€/GJ] 12.66 9.48 9.39 9.30

4.5.15 HTC and subsequent combustion in a biomass fired CHP plant

In section 4.5.1, it was shown that pretreatment with HTC can potentially make com-bustion more efficient for biomass with a high moisture content, because the amount ofwater evaporated from the feedstock during drying or combustion is reduced. This sectionanalyzes the impact of HTC pretreatment on the overall conversion chain efficiency inmore detail, taking into account the efficiency and the auxiliary energy consumption ofthe HTC plant models.

It is assumed that the biocoal is fired in a medium-scale biomass-fired CHP plant. Utilizingthe biocoal in a plant being fired exclusively with biomass may be advantageous in someEuropean countries because it may then qualify for a renewable energy feed-in tariff. Theconversion chain efficiency of HTC plant and subsequent CHP plant is compared to thedirect utilization of the raw biomass in a CHP plant. To this end, simulations of theCHP plant with a fluidized-bed boiler are performed, firstly fuelled with the raw biomass,namely wood chips and PGW-70, and secondly with the respective biocoals obtained fromboth feedstocks in the HTC plant simulations HTC-1.00 and HTC-3.00.

All simulation cases deliver 32 MW of thermal energy. Steam turbine inlet parametersare 500°C and 80 bar. For the CHP plants processing wood, simulation models comprisethe combustion of wood-derived biocoal, combustion of the raw wood and combustion ofwood pre-dried to 10% with low temperature heat taken from the CHP plant. For theplants processing PGW the feedstock is dried to a water content of 50% or lower in allsimulation cases to enable a stable combustion. An additional case with SSD drying allowsthe benefits of utilizing the heat of condensation to be analyzed.

Two scenarios regarding the heat demand are considered: First, the production of processsteam at 10 bar, and second, the production of hot water at 95°C for district heating. Theflowsheet of both scenarios (including pre-drying of the feedstock) is shown in Figure 4.14.Water returns at 150°C and 5 bar in the case of process steam production and at 55°C fordistrict heating (36).

The analyzed simulation cases are listed in Table 4.34. The term LTD-10% indicates lowtemperature drying to a water content of 10%.

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The flowsheet of simulation case SC-3.1.4 with SSD is shown in section B.4.15. Heatingsteam is provided to the drier from a steam turbine extraction at 15 bar. Thermal energyis recovered from the drier exhaust steam and heating steam condensate for raising steamat 4.5 bar, which is then expanded in the low pressure steam turbine stage. A similarconfiguration is employed in a combined heat, power and wood pellets trigeneration plantin Sweden since 1997 [324].

The results of the energy and exergy analyses are summarized in Table 4.35. Details forall simulation cases are given in section B.4.15.

steamturbine

exhaustgas

exhaustgas

dried biomass

air

airash

5

13

311416

1519 29

28

18

2

17

35

38 37 3627

9

6

10

drierW1

W5

W2

W3

G

wetbiomass

steamturbine

exhaustgas

exhaustgas

dried biomass

air

airash

5

13

31

1424

25

16

15

1918

2

17a) b)

32

30 31

9

6

10

drierW1

W5

W2

W3

G

Figure 4.14: Flowsheets of the biomass fired CHP plants supplying process steam (a) andhot water for district heating (b).

Table 4.34: Simulation cases for HTC plants with subsequent combustion and for thecombustion of raw or dried biomass.

feedstock heat demand upgrading dryingSC-1.0.1 wood steam HTC —SC-1.0.2 wood steam — —SC-1.0.3 wood steam — LTD-10%SC-1.1.1 wood hot water HTC —SC-1.1.2 wood hot water — —SC-3.0.1 PGW steam HTC —SC-3.0.2 PGW steam — LTD-50%SC-3.0.3 PGW steam — LTD-10%SC-3.1.1 PGW hot water HTC —SC-3.1.2 PGW hot water — LTD-50%SC-3.1.4 PGW hot water — SSD-10%

The electrical efficiency of all analyzed systems is a relatively low 8–17%, due to the lowsteam parameters and the back-pressure operation of the steam turbine. The electricalefficiency of the CHP plant itself ranges between 10–21% based on HHV, or 13–27% basedon LHV.

Generally, the systems delivering hot water for district heating have higher electrical andenergetic efficiencies than comparable systems producing process steam. The exergeticefficiencies however are lower because of the low specific exergy of the hot water.

Drying increases the electricity output and electrical efficiency, because the drier provides

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Table 4.35: Conversion chain efficiency for HTC with subsequent combustion and for thecombustion of raw or dried biomass.

rawbiomass

electricity(net)

steam /hot water

electricalefficiency(HHV)

energeticefficiency(HHV)

exergeticefficiency

[MWHHV] [MWel] [MWth] [–] [–] [–]

SC-1.0.1 55.40 4.09 31.98 7.4% 65.1% 27.0%SC-1.0.2 54.76 5.92 31.98 10.8% 69.2% 30.5%SC-1.0.3 60.12 8.36 31.98 13.9% 67.1% 31.4%SC-1.1.1 58.14 7.75 31.90 13.3% 68.2% 21.5%SC-1.1.2 54.30 9.34 31.90 17.2% 76.0% 25.9%SC-3.0.1 58.91 3.36 31.98 5.7% 60.0% 23.9%SC-3.0.2 108.13 13.58 31.98 12.6% 42.1% 21.6%SC-3.0.3 155.47 24.06 31.98 15.5% 36.0% 21.0%SC-3.1.1 61.68 6.99 31.90 11.3% 63.0% 18.9%SC-3.1.2 105.50 15.80 31.90 15.0% 45.2% 18.6%SC-3.1.4 48.88 7.78 31.90 15.9% 81.2% 25.4%

an additional heat sink which allows more electricity to be produced in cogeneration.However, drying also increases the fuel demand and decreases the energetic efficiency.

Pretreatment with HTC decreases the electrical, energetic and exergetic efficiencies by4–8 percentage points when wood is used as the feedstock. When PGW-70 is used as afeedstock, HTC increases the energetic efficiency by 18 percentage points and the exergeticefficiency by 4–7 percentage points. The electric efficiency however is decreased by HTCby 1–4 percentage points. The best performance is achieved under SSD drying, whichincreases the energetic efficiency by 36 percentage points and the exergetic efficiency by 7percentage points.

The exergy analysis reveals that the combustion has a dominating effect on the perform-ance of the overall system: 54–67% of the exergetic fuel provided to the CHP plant isdestroyed in the boiler, while all the remaining CHP components together destroy only4–16%. The exergetic efficiency of the boiler is 7–8 percentage points higher when biocoalis combusted compared to biomass with a water content of 50%. The increased efficiencyof the boiler plus the avoided exergy destruction of the drier outweigh the conversion lossesof the HTC process when PGW-70 is used as the feedstock.

The energetic efficiency ranges between 65–76% for systems processing wood. Configura-tions employing HTC perform worse than the combustion of the raw or dried wood, fromboth an energetic and exergetic perspective. For systems processing PGW the energeticefficiency is as low as 36–45% when conventional drying is employed. This clearly makescombustion of wet feedstocks unattractive. With HTC, the energetic efficiency can beincreased to 60–63%. Superheated steam drying with recovery of the heat of condensationfrom the fuel moisture performs even better than HTC, with an energetic efficiency of81%. However, this requires a low temperature heat sink, such as district heating at arelatively low temperature level. The exergetic efficiencies of the analyzed systems rangebetween 18–31%. Of all the analyzed cases processing PGW, cases SC-3.1.4 with SSD

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and SC-3.0.1 with HTC have the highest exergetic efficiencies with 25.4% and 23.9%,respectively.

The analysis confirms that despite the conversion losses of the HTC process, HTC canimprove the conversion chain efficiency, but only for very wet feedstocks.

The easier handling and higher energy density of the HTC biocoal compared to raw bio-mass may result in some savings in investment and operating costs for the boiler andfeedstock preparation equipment. Due to the higher calorific value of the HTC biocoal,the flue gas volume per MJ of thermal boiler load is reduced by more than 30% com-pared to the combustion of raw wood. These potential savings and the gains in efficiency,however, are unlikely to offset the more than twofold increase in fuel cost induced by theHTC treatment. For example, under the assumptions employed in this work, the levelizedcost of raw wood (SR) including transport and storage amounts to 5.4 €/GJ, whereas thebiocoal cost from HTC-1.00-m is 14.3 €/GJ.

4.6 Comparison of the biomass upgrading processes

In this section, the different upgrading processes are compared. The focus lies on thetechnical feasibility, economic viability and GHG mitigation cost of replacing fossil fuelsin existing power stations.

4.6.1 Exergetic performance of the biofuel production

Figure 4.15 summarizes the results of the exergy analysis for the various biofuel productionprocesses. The first three cases WP-1.0, TOR-1.0 and HTC-1.0 process wood, while theremaining four process wet biomass, namely organic waste or grass silage.

The exergy destruction of the reactor converting biomass into biofuel is much higher forthe anaerobic digestion than for the processes producing solid biofuels. This is not sosurprising, because it is a deeper conversion. The feedstock is decomposed into a gas,while in torrefaction and HTC, in essence, just a few functional groups are removed fromthe biomass.

The hydrothermal processes suffer relatively large exergy losses due to the residue streamscontaining organic compounds that are not converted to biofuel. For HTC, the mainexergy loss is the waste water with dissolved organic compounds. For anaerobic digestion,it is the hard-to-digest part of the feedstock, which leaves the process as digestate. ForADM-3.0 with a low methane yield, this results in a biomass consumption twice as highas for HTC. Even for ADM-3.1 with optimistic assumptions regarding the methane yield,the biomass consumption is higher than for the processes producing solid biofuels due tothe exergy destruction in the reactor and the exergy loss with the digestate. Dependingon the availability of waste biomass or the land for bioenergy crop production, efficientutilization of the biomass resources may well become an important criteria.

Drying of the solid biofuel and the combustion to provide thermal energy for the dryingprocess represents the biggest source of exergy destruction in all processes producing solid

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0.0

0.2

0.40.6

0.8

1.0

1.21.4

1.6

1.8

2.0

2.22.4

2.6

2.8

WP-1.0

TOR-1.0

HTC-1.00

HTC-3.00

ADP-3.0

ADM-3.0

ADM-3.1

[MJe

x ]

biomass electricity, consumption

biofuel electricity, productionsolid / liquid residue other waste streams

ED reactor ED combustion + dryingED other

left bar

right bar

from wood from wet feedstocks

Figure 4.15: Exergy of biomass and electricity consumption, waste streams and mainsources of exergy destruction for selected biofuel plant models, normalized to 1 MJ ofupgraded biofuel. For each case, the left bar signifies the inputs and the right barsignifies the outputs and exergy destruction.

biofuels. In HTC, this exergy destruction is about halved by the facilitation of mechanicaldewatering. On the other hand, exergy destruction within miscellaneous plant componentsis higher in HTC due to the more complex process.

ADP-3.0 has a higher biomass consumption and exergy destruction than the other pro-cesses for solid biofuel production, but delivers electricity as a byproduct. Its performancecannot therefore be directly compared to the other processes, which deliver biofuel as thesole product.

4.6.2 Conversion chain efficiency for the combustion of upgraded biofuels inexisting power stations

The most widely discussed utilization of torrefied wood pellets and HTC biocoal pellets isco-combustion in coal-fired power stations, which are generally more efficient than typicalbiomass-fired plants. The electrical efficiency of a state of the art bituminous pulverizedcoal-fired power station is about 43% (LHV), while the average efficiency of the bituminouscoal-fired power station operating in Germany is 38% (LHV) [341]. Substituting fossil fuel

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in existing power stations by biofuels offers power plant operators the flexibility to reactto rising CO2 prices and may provide a short term option to increase the use of bioenergy.

Table 4.36: Efficiencies of wood pellets, biocoal, bituminous coal and biomethane com-bustion in conventional power stations.

upgrading WP

1.0

HTC

1.00

HTC

3.00

HTC

4.00

TOR

1.0

ADP

3.0

ADM

3.0

ADM

3.1

feedstock wood wood PGW MOW wood grass coal grass grass

HHV efficiencies

power plant 36.3% 37.7% 37.8% 37.2% 36.7% 36.1% 38.3% 49.6% 49.6%

conversion chain 28.7% 27.1% 24.3% 21.8% 28.3% 29.1% 17.1% 32.9%

LHV efficiencies

power plant 39.7% 39.9% 39.9% 39.4% 39.8% 39.7% 40.0% 55.0% 55.0%


exergetic efficiencies

power plant 34.5% 36.3% 36.3% 35.7% 35.1% 34.0% 37.2%


capacity loss1) 38% 12% 18% 39% 31% 45%1) at 100% biofuel

The conversion chain efficiencies from raw biomass to electricity are calculated as explainedin section 3.2.5. The results are summarized in Table 4.36. The energetic efficiency of thereference plant fired with bituminous coal is 40% (LHV), or 38.3% (HHV). Compared tobituminous coal, the penalty on the HHV-based efficiency of the power plant ranges from0.4 percentage points for HTC biocoal from wood to 2.2 percentage points for ADP pelletsfrom grass press cake. The LHV-based efficiency is less affected by the fuel composition,the penalty ranging from 0.1–0.3 percentage points.

Due to the lower calorific value of the biofuels compared to bituminous coal, the capacityof the power station is reduced, given that the fuel mass flow remains unchanged. Whenthe power station is fuelled entirely by biofuel, the capacity reduction ranges from 12–45%for the biofuels considered in this work. This may lead to an additional efficiency dropbecause the steam cycle is operated in part load. For co-firing, the capacity and efficiencyloss will depend on the coal to biofuel ratio, but also on the set-up of the individual powerstation. These effects are not considered in this work. At Tilbury power station, UK,conversion to 100% wood pellets reduced the capacity by 30% and the efficiency by 1.7percentage points [109]. For torrefied wood and HTC biocoal, the efficiency loss shouldbe lower because their calorific value is higher.

Taking into account the upstream conversion losses associated with the biomass upgrad-ing, the HHV-based efficiencies using solid biofuels range between 22–29%. Using woodas a feedstock, the conversion chain efficiency is highest when wood is simply pelletized.Compared to pelletizing, pretreatment with torrefaction and HTC reduces the HHV effi-ciency by 0.4 and 1.6 percentage points, respectively. The LHV-based efficiency is higherfor the overall conversion chain than for the power plant in some cases where feedstockswith 70% water content are treated in an upgrading process which removes water in liquidstate (compare section 3.6).

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Given the relatively minor differences in efficiencies for the various pretreatments technolo-gies, other considerations may be decisive for the selection of the most suitable upgradingtechnology. HTC biocoal is the most coal-like biofuel, with the highest calorific value andprobably the best properties regarding storage and feed preparation for pulverized coalburners. The properties of torrefied wood lie somewhere in between wood pellets andHTC biocoal. There is no information available on the quality of the pellets producedfrom press-cake (ADP), but they are likely to be similar to wood pellets at best.

The exergetic efficiencies for the conversion chains with solid biofuels lie between 20–27%, in the same range as for the utilization of raw biomass or biocoal in the CHP plantdescribed in section 4.5.15.

Since biomethane is upgraded to natural gas quality for feed-in to the natural gas grid,there is no capacity or efficiency loss at the receiving combined cycle power station. Theconversion chain efficiency (HHV) is strongly dependent on the methane yield, 17.1% forADM-3.0 and 32.9% for ADM-3.1. The conversion pathways employing a solid upgradedbiofuel and its utilization in a pulverized coal-fired power station offer a more efficientconversion from biomass to electricity, unless a very high methane yield is achieved withanaerobic digestion.


If no special incentives are available to support co-firing biomass, the upgraded biofuelshave to compete economically with the fossil fuel they seek to displace plus the ETScarbon certificates required for the combustion of those fossil fuels. With CO2 pricesbelow 15 €/t, it is difficult for biomass to compete with fossil fuels. Therefore, biofuels aremostly utilized in applications where they receive additional support by feed-in tariffs orrenewable certificate schemes. Due to such schemes, the traded price for biofuels can besignificantly higher than the fossil fuel price plus carbon certificates. Therefore, it is morelikely for upgraded biofuels to make their entry in these policy-mandated biofuel marketsthan in direct competition with fossil fuels.

It can be reasonably assumed that the upgraded solid biofuels under consideration areeconomically viable if their production and delivery costs remain below the wood pelletmarket price.

For the biomethane cases ADM-3.0 and ADM-3.1, the biomethane market price wouldact as the benchmark. This market prices depend on country-specific feed-in tariffs, taxes,renewables certificates schemes and feed-in regulations for biomethane. In Germany, bio-methane can be injected into the natural gas grid when upgraded to the required quality.Biomethane is mostly purchased for utilization in small CHP plants, because the electri-city produced then qualifies for a feed-in tariff under to the Renewable Energy SourcesAct (Erneuerbare Energien Gesetz — EEG) [282]. According to a survey amongst biogastraders, the average price of biomethane lay between 65 and 83 €/MWh for the period2008–2010 [219].

The use of industrial wood pellets in large facilities is common in Belgium, the Netherlands,UK and Sweden, because their legislation, in contrast to the German feed-in law, supportsbiomass co-firing [10, 48]. The various support schemes allow utility operators to pay wood

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pellet prices of 100 €/t (UK) to 160 €/t (Sweden) [10]. However, some of these countriesplan to reduce their support for the co-firing of biomass in large power stations in the nextyears and focus instead on small-scale electricity production fired exclusively on biomass[48].

Figure 4.16 shows the sum of fuel plus carbon certificate costs in relation to the CO2price for bituminous coal, pulverized lignite22, raw lignite and natural gas. Market pricesfor wood pellets (EU) and biomethane (Germany) are also displayed. Of the fossil coals,pulverized lignite is the most expensive, because it is a preprocessed fuel. At CO2 pricesbetween 15–30 €/t, pulverized lignite becomes more expensive than wood pellets. Sub-stituting pulverized lignite for biocoal in industrial drying or CHP facilities would seemmore economically attractive than substituting unprocessed fossil coal in large coal-firedpower stations. For plant operators, the possibility to switch to biocoal may provide aninteresting option for reducing the cost risk related to rising CO2 prices.

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WP, 115 €/t

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Figure 4.16: Levelized fuel plus carbon certificate costs in relation to the CO2 price.

Unless otherwise stated, the results presented in the following are based on the assumptionthat solid biofuels substitute bituminous coal in a pulverized coal-fired power station,and biomethane replaces natural gas in a combined cycle plant. The economic viabilityin respect to prevailing wood pellets and biomethane market prices, and substitutingpulverized lignite by HTC biocoal, are also discussed.

Figure 4.17 and Figure 4.18 show the production costs and achievable revenues for up-graded biofuels from wood, waste and grass. For every upgrading scenario, the left columnshows the revenues (rev.), which can be achieved when competing with fossil fuel plus car-bon certificates with a CO2 price of 15 €/t (FOS). The second-to-left column shows therevenues which can be achieved in the European market for industrial wood pellets, or

22Pulverized lignite is preprocessed by drying and milling and is used as a fuel for industrial drying andCHP facilities. See section 2.2.3.11.

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the biomethane market in Germany (RE), assuming market prices of 130 €/t for woodpellets and 75 €/MWhHHV for biomethane. The 1–6 columns on the right side show theproduction cost of the biofuel, including delivery to the power station, for various plantcapacities (s, m, l), and, in the case of wood, various feedstocks (SR, FR).

02

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WP-1.0 TOR-1.0 HTC-1.00

cost

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iofu

el [€

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HHV ]

feedstock feedstock transportbiofuel road transport to harbour biofuel shippingbiofuel transport, regional carrying chargeslabour electricityother operating cost ash disposalwaste disposal co-produced electricityfossil fuel replacement wood pellets / biomethaneavoided ETS certificates avoided GHG, waste

Figure 4.17: Biofuels from wood: revenues and production cost (at power plant gate) perunit of biofuel. Red borders indicate revenues, black borders indicate costs.

HTC-4.00-m is the only scenario in which the biofuel is economically viable against fossilfuels. This is because of the high remuneration for the MOW disposal, which completelycovers the production cost without requiring revenues from fossil fuel replacement andavoided carbon certificates. Medium to large scale wood pellet plants using forest residuesin North America and importing the pellets (WP-1.0-l-FR), and anaerobic digestion fromgrass with high yield (ADM-3.1 ) are economically viable against the wood pellet andbiomethane market prices. HTC from EFB in case HTC-5.00-m* can compete in thewood pellet market if the avoided emissions from EFB dumping can be monetized. HTC-3.00-m comes close to being economically viable, its th biofuel costs being 8% higher thanthe revenues from the wood pellet market price.

Wood pellets from domestic short rotation coppice are 25–35% more expensive than woodpellets produced overseas due to the high cost of feedstock. They are also 7% moreexpensive than HTC biocoal from HTC-3.00-m. HTC from waste therefore seems aninteresting option if the use of domestic biomass sources is to be expanded.

Torrefied wood pellets are 0.0–0.3 €/GJ more expensive than conventional wood pellets.However, they are more coal-like and therefore better suited for power stations designed for

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feedstock feedstock transportbiofuel road transport to harbour biofuel shippingbiofuel transport, regional carrying chargeslabour electricityother operating cost ash disposalwaste disposal co-produced electricityfossil fuel replacement wood pellets / biomethaneavoided ETS certificates avoided GHG, waste

Figure 4.18: Biofuels from waste and grass: revenues and production costs (at power plantgate) per unit of biofuel. Red borders indicate revenues, black borders indicate costs.

coal. Modifications to facilitate biomass co-firing at a coal-fired power station may requirean investment of 50–300 $/kWbiomass [11]. A value of 55 €/kWbiomass for enabling the co-firing of wood pellets would translate to an additional 0.3 €/GJbiofuel. If the power plantmodifications can be avoided through the use of torrefied pellets, the higher productioncost of torrefied pellets may be compensated.

Since torrefied pellets are more water resistant then conventional wood pellets, they canpotentially be stored outside, which could lead to further cost reductions for transport,storage and handling at the harbour. In the scenarios where pellets are produced overseasand shipped to Europe, biofuel transport contributes to up to 60% of the total biofuelcost. Differences in the transport and storage modalities may therefore be the decidingfactor regarding the relative cost competitiveness of imported torrefied and conventionalwood pellets. Specific supply chains need to be analyzed in more detail than is done inthis work to reveal whether torrefaction could be worthwhile.

The ship transport accounts for 16–19% of the overall pellet costs in the scenarios em-ploying overseas forest residues or EFB as the feedstock. As discussed in section 3.2.3.2,shipping prices depend on whether a well established trade route is used and fluctuatestrongly in response to the supply and demand of shipping capacity. The development ofthe trade in dry bulk goods on the respective shipping routes may therefore prove crucialto the economic viability of HTC biocoal from EFB and imported conventional or torrefied

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wood pellets.

HTC biocoal from wood is 50–85% more expensive than conventional wood pellets andtorrefied wood, due to the higher investment and operating costs of the notibly morecomplex process.

For HTC-3.00 and HTC-4.00, the revenues received from PGW or MOW disposal areequivalent to or higher than the revenues from selling the biocoal. In this regard, HTCshould be seen as primarily a waste treatment technology rather than a biofuel productionprocess. For a comprehensive analysis of its market potential, HTC needs to be comparedwith other waste treatment technologies such as composting and incineration.

The production of biofuels from grass is expensive for all the technologies considered, dueto the high feedstock cost and small plant capacities. The feedstock costs alone are higherthan the achievable revenues from substituting fossil fuel or from the wood pellet marketprice. Due to the high market price for biomethane, case ADM-3.1 is economically viablenevertheless. The biomethane yield is a major determinant for the cost competitiveness ofthe anaerobic digestion scenarios. ADM-3.0 with a low yield is one of the worst performingupgrading scenarios, and ADM-3.1 with a high yield is one of the best.

Upgrading technologies may facilitate the large-scale use of biomass in existing power sta-tions without requiring investment in power plant modifications. However, the investmentrequired to build the upgrading plants is considerable. If the upgraded fuels are burnedin existing power stations as described above, the required investment amounts to 290–830 €/kWel for wood pelletizing, 480–1130 €/kWel for torrefaction, 2200–5900 €/kWel forHTC, 2730 €/kWel for ADP and 2940–6000 €/kWel for ADM. For comparison, the specificinvestment for dedicated wood-fired power stations is reported in the range of 1360–6750€/kWel [244, 245, 247, 342]. Adjusting an existing coal-fired power station to biomassco-firing requires an investment of approximately 170–1000 €/kWel [11].

4.6.4 GHG balance and mitigation cost

The GHG emissions arising from biofuel production and the avoided emissions by repla-cing fossil fuel at a power station are summarized for selected scenarios in Table 4.37. TheGHG reduction per unit of feedstock biomass amounts to 55–77 kg/GJbiomass for all scen-arios employing solid upgraded biofuels except for HTC-5.00, where the avoided methaneemissions from dumping EFB lead to a GHG reduction of 226 g/GJbiomass.

For anaerobic digestion, the GHG reduction is 15–30 kg/GJbiomass, considerably lowerthan that for solid biofuels. Since natural gas is replaced (instead of bituminous coal),the avoided emissions per GJ of biofuel are 40% lower. Electricity consumption for thebiogas compression and fugitive methane emissions are responsible for relatively high GHGemissions during the production of the biomethane. Avoided GHG emissions by replacingmineral fertilizer by digestate are not considered in this analysis and may improve theGHG balance of anaerobic digestion.

For the WP, TOR and HTC scenarios, the GHG emissions from using biofuels amountto 5–22% of those from using fossil coal in a power station. For ADM-3.0 and ADM-3.1, the GHG emissions amount to 35–41% of those using natural gas. GHG emissionsfrom anaerobic digestion reported in literature vary widely. Meyer-Aurich et al. report

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that GHG emissions estimates from various studies lie in the range of 7–90% relative tothe fossil fuel reference system [45]. This illustrates the importance of the biomethanesupply chain for the GHG reduction and for the GHG mitigation cost. A more detailedanalysis than offered in this work, including different scenarios for methane losses and theutilization of the digestate, is required for a comprehensive assessment.

Table 4.37: GHG emissions related to biofuels and displaced fossil fuels per unit of biofuelin [kgCO2,eq/GJHHV].

WP1.0-m

SR

WP1.0-l

FR

TOR1.0-l

FR

HTC

3.00-m

HTC

3.90-s

HTC

5.00-m

ADP

3.0-s

ADM

3.0-s

ADM

3.1-s

biofuel

production1) 4.6 12.6 10.3 21.2 10.1 -193.6 -12.6 23.3 19.8

transport1) 1.9 1.9 1.6 1.5 1.5 1.3 2.2 0.7 0.7

total 6.5 14.6 11.9 22.7 11.6 -192.3 -10.5 23.9 20.5

avoided emissions by fossil fuel replacement

supply chain 11.7 11.7 12.0 12.2 12.2 12.1 11.6 7.7 7.7

combustion 86.3 86.3 88.1 89.9 89.9 88.9 85.6 50.5 50.5

total 98.1 98.1 100.1 102.1 102.1 101.0 97.3 58.2 58.2

net GHG reduction

per GJ biofuel 91.5 83.5 88.1 79.4 90.5 293.3 107.8 34.3 37.7

per GJ feedstock 76.7 70.0 70.7 60.1 73.2 226.0 70.1 14.7 29.91)For the biofuels produced overseas, the emissions listed here for biofuel production include delivery to a port in

Europe. The emissions for transport comprise 100 km truck transport from either the harbour or the biofuel plant

to the power station for the solid biofuels, and 100 km pipeline transport for biomethane.

Figure 4.19 shows the GHG mitigation costs for selected scenarios in relation to the GHGreduction per unit of feedstock biomass. The results of all the cases analyzed are sum-marized in Table B.43.

The GHG mitigation costs for conversion pathways with HTC from waste amount to -25€/t to 127 €/t, depending on HTC plant capacity and remuneration for waste disposal.For a medium-scale HTC plant processing PGW with a remuneration of 20 €/tFM, theyamount to 77 €/t, similar to a medium-scale wood pelletizing plant using short rotationwood, at 75 €/t. Large-scale wood pelletizing plants using forest residues in North Americaand importing the pellets to Europe gives 53 €/t.

When HTC biocoal from HTC-3.00-m substitutes pulverized lignite rather than bitumin-ous coal, the GHG mitigation costs drop to 47 €/t.

Torrefaction results in 1–5% higher GHG mitigation cost than wood pelletizing with thesame plant capacity and feedstock. HTC from wood yields significantly higher costs thantorrefaction and wood pelletizing, namely 129 €/t for a medium-scale plant using SRwood.

HTC from EFB offers a very high GHG reduction per unit biomass at a moderate costof 28–34 €/t, because of the avoided methane emissions from EFB dumping. However,any treatment method replacing EFB dumping would lead to a similarly high emission

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Figure 4.19: GHG mitigation cost and GHG reduction per unit feedstock biomass. Theright diagram is a magnification of the boxed area in the left diagram.

reduction. Therefore, alternative technologies such as composting, anaerobic digestion oreven just mulching the plantation should be analyzed as well and compared to HTC.

The anaerobic digestion of grass leads to a low GHG reduction per unit of biomass atvery high costs of 398–917 €/t. These costs are partly due to the high feedstock cost, butthe production of solid biofuels from grass with ADP or HTC leads to significantly lowercosts of 170 and 261 €/t, respectively. The influence of the feedstock cost on the GHGmitigation costs of ADM, ADP and HTC are discussed in section 4.6.6.

GHG mitigation costs for anaerobic digestion from energy crops (typically maize) reportedin literature range from 200–520 €/t without considering the effects of land use change[229, 334, 343]. However, they cannot be directly compared to the values obtained in thiswork because they usually refer to the displacement of average electricity from the Germanpower plant fleet by new small-scale CHP plants fired with biogas. Mitigation costs below300 €/t are mostly obtained when credits for co-produced thermal energy from CHP aretaken into account. Exclusive electricity production leads to mitigation costs of 378–520€/t [229, 343]. Case ADM-3.1 which has a biomethane yield comparable to maize resultsin mitigation costs at the lower end of this spectrum.

4.6.5 Identifying the most suitable plant size

For a new technology like HTC, the question of suitable scale arises. While traditionally,small decentral plants were advocated for biomass conversion, the recent developments inwood pelletizing show a trend to very large processing capacities of over 500 MWbiomass[110]. As can be seen in Figure 4.18, a larger plant capacity leads to a decrease in carryingcharges and labour cost per GJ biofuel due to economy-of-scale effects. On the other handthe feedstock transportation cost increases, because biomass has to be collected from alarger area. In this section, the influence of plant capacity on the respective cost positions

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is analyzed with more detail for HTC and WP in order to explore the optimal plant size.Subsequently, the biomass supply situation is discussed.

Based on the results of the cost estimates for small, medium and large plants, scalingexponents according to Equation 3.11 for the total capital investment (TCI) of the biofuelplants are derived. Figure 4.20 shows the specific investment per unit of product in relationto the feedstock input capacity on a semi-logarithmic scale. Investment cost for ADM andADP, for which cost estimates have only been performed for one plant size, are alsodisplayed. Cost data from literature for wood pelletizing [106, 110, 112, 326], torrefaction[114, 122, 139] and anaerobic digestion [218, 334, 344] are included for comparison.

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Figure 4.20: Specific investment costs of biofuel production plants in relation to the plantcapacity. The x-axis is logarithmic. Literature data from [106, 110, 112, 114, 122, 139,218, 326, 334, 344].

It becomes obvious that all the technologies suitable for wet biomass like grass and waste,namely HTC, ADM and ADP, have considerably higher specific investment costs thanwood pelletizing and torrefaction. The investment cost of HTC is subject to a strongeconomy-of-scale effect.

Table 4.38 shows the cost for 100 km road transport by truck for the various feedstocksand upgraded biofuels. The costs are given per GJLHV because the LHV better representsthe usable energy at a power station. The transport cost are reduced by 73–93% whenbiocoal pellets are transported instead of raw biomass. While the transport cost for woodchips are moderate, those of waste can be as high as 7.5 €/GJ. This shows that the locationof the upgrading plant should be close to the source of biomass and that long distancetransport of wet feedstock material should be avoided. This may limit the capacities ofbiofuel plants processing wet waste biomass.

When HTC biocoal is used as a substitute for pulverized lignite, it needs to be transportedin pulverized form. This is more than twice as expensive as the transport as pellets, butstill appears acceptable at less than 1 €/GJ.

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Table 4.38: Biomass and biocoal road transport costs for a 100 km distance.

raw pelletized HTC pellets HTC pulverizedwood chips [€/GJLHV] 1.29 0.50 0.35 0.91

PGW-70, grass [€/GJLHV] 4.28 0.56 1) 0.38 0.99MOW [€/GJLHV] 7.49 0.51

PGW-50 [€/GJLHV] 3.25 0.391) ADP pellets

Figure 4.21 shows a breakdown of the biofuel production costs for HTC-3.00 in relationto the plant capacity. The feedstock transport costs increase significantly with the plantcapacity, but are overcompensated by the decrease in carrying charges and labour cost forplant capacities of up to 200 kt/a. In fact, there is a sharp decrease in the overall biocoalproduction costs up to about 100 kt/a. Beyond that, the total production costs show littlechange. A doubling of capacity from 10 to 20 kt/a leads to a decrease in specific biocoalcosts by 9.6 €/GJ, while doubling the capacity from 100 to 200 kt/a only reduces biocoalcosts by 1.2 €/GJ.

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Figure 4.21: Break-down of biocoal cost versus plant capacity for HTC-3.00.

At most locations, there will be only a limited amount of biomass available in the sur-roundings. With rising plant capacity it will doubtlessly become increasingly harder toprocure the required amount of biomass at a reasonable distance. For example, for an HTCplant processing organic waste or park and gardening waste, the required feedstock maybe available with a high yield per hectare within a densely populated city area, whereasin the surrounding countryside the yield per hectare will normally be much lower.

Figure 4.22 shows the biofuel production cost and GHG mitigation cost of wood pel-

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lets from short rotation and HTC biocoal from PGW-70 for different feedstock supplyscenarios. The continuous lines represent the base cases, and the dashed lines representscenarios with a poorer feedstock supply situation. The yield per hectare is reduced to10% of the base case for SR wood and 50% of the base case for PGW.

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Figure 4.22: Biofuel production cost and GHG mitigation cost for different feedstock sup-ply scenarios.

The WP production costs are not very sensitive to the feedstock supply situation. Forthe low supply scenario, where only 1% instead of 10% of the surrounding land area isproviding wood to the biofuel plant, the wood pellet production costs of a plant handling300 kt/a are just 0.63 €/GJ (7%) higher than in the base case.

For HTC from PGW, the situation is different. Due to the higher transportation cost, thelower yield per hectare and consequent longer transport distances, the production costsare much more sensitive to the feedstock supply situation. For a plant processing 300kt/a, the low supply scenario results in a biocoal cost increase of 2.0 €/GJ (25%). For a100 kt/a plant the difference is still 1.0 €/GJ (11%).

The negative effect of a lower biomass availability is even more pronounced for the GHGmitigation costs, where higher transportation costs are accompanied by higher GHG emis-sions from the transport. For a 300 kt/a plant, this leads to a difference of 33 €/t (45%)between the two supply scenarios.

Depending on the feedstock availability, the cost optimal feedstock processing capacity liesin the range of 250–900 kt/a (100–350 MW) for WP. This is in good agreement with thecapacity of plants built recently in North America. For HTC from waste, the cost optimalfeedstock capacity is 150–250 kt/a (30–50 MW) for the two supply scenarios analyzedhere. The availability of waste feedstocks is likely to be highly site-specific and mustbe carefully analyzed in order to assess the feasibility of an HTC project. The amountof waste available for a new HTC plant not only depends on the total amount of wastegenerated in the area, but also on existing long term contracts for waste processing atother facilities (such as composting).

The results suggest that a feedstock processing capacity of around 100 kt/a is required tomake HTC biocoal from PGW cost competitive with wood pellets. This is roughly the

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equivalent of the annual park and gardening wastes from 2 million people. Average com-posting works process 10–20 kt/a [280]. This suggests that locations with a sufficientlylarge feedstock supply are likely to be rare. For example, the sum of separately collec-ted organic household waste, dead leaves, park and gardening waste, fruit and vegetableresidues and food waste from canteens, restaurants and hospitals in Berlin amounts to242 kt/a (see Table 2.4). Dead leaves, which accrue only in autumn and may produce lowquality biocoal due to their high ash content, account for 40% of this potential.

Waste streams from biomass processing industries may be among the most promisingfeedstocks for HTC since large amounts of biomass are potentially available year-roundat one site. A palm oil mill, for example, delivers 20–90 kt/a EFB (see section 2.1.2.5).Palm oil production and other significant biomass processing industries should be furtherinvestigated to identify potential feedstocks for HTC.

HTC can offer a niche role for waste processing, particularly for wet feedstocks. Evenat sites with a low waste availability, HTC may still be economically viable comparedto other treatment options, since the costs for alternative waste treatment technologiesequally increase with decreasing plant size. This should result in a higher remunerationpaid for waste treatment. As discussed in section 4.6.3, the waste disposal revenues are themost important source of income for HTC plants. Remuneration rates for waste treatmentvary widely from 20–100 €/t for MOW and 15–45 €/t for green waste [271, 345]. Witha remuneration of 45 €/t for green waste, an HTC plant capacity of 25 kt/a leads toproduction cost of 10.78 €/GJ, equal to that of a 100 kt/a plant with a remuneration of20 €/t.

4.6.6 Solid biofuels versus biomethane

An interesting question is whether a feedstock suitable for both HTC and anaerobic di-gestion should be converted to solid biofuel pellets or to biomethane. Biomethane can beused to displace natural gas rather than coal, a more expensive, higher quality fuel, whichcan be converted to electricity at a higher efficiency.

In the results presented so far, HTC was mostly discussed using PGW as a feedstock,while grass silage from extensive cropping served as the feedstock for anaerobic digestion.Grass is the most expensive feedstock considered in this work, while processing PGWyields revenues for waste disposal. This makes HTC, ADM and ADP difficult to compare.

In principle, HTC, ADP and ADM are all capable of processing various types of wetnon-woody biomass, including grass silage and waste biomass. The anaerobic digestionprocess needs to be carefully tailored to the feedstock, and the yield depends stronglyon the feedstock composition. For example, the energy yield from grass of the samespecies may vary between 15% and 70% depending on the time of harvest [221, page 68].Feedstocks such as roadside grass which contain heavy metals and polycyclic aromatichydrocarbons are banned for anaerobic digestion because the pollutants could be spreadon agricultural soil with the digestate [57]. HTC, on the other hand, is tolerant in respectto feedstock composition, although a high mineral matter content may accumulate in thebiocoal and decrease its quality as a power plant fuel. No work on the ADP process forvarious feedstocks is known to the author, but it is likely to lie in between ADM and HTCregarding its tolerance to feedstock quality.

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In this section, the feedstock price is varied ranging from −20 €/t to 30 €/t for HTC,ADP and ADM. Energy yields and all costs except the feedstock cost are kept constantfor the analysis. In reality, feedstock obtained at very low or negative cost is likely tobe of lower quality, resulting in lower biomethane yields and higher processing costs foranaerobic digestion.

Since ADM produces a gaseous fuel and HTC and ADP produce solid fuels, it makeslittle sense to compare the cost of the biofuels directly. Therefore, a different metricis employed. The additional cost Δc per unit of electricity relative to fossil fuel use isanalyzed, assuming that biomethane replaces natural gas in a combined cycle plant andbiofuel pellets replace bituminous coal in a pulverized coal-fired power station. Hence:

Δc = cbiofuelηfossil

ηbiofuel− cfossil − cCO2fCO2,fossil (4.6)

where cbiofuel and cfossil are the cost of the biofuel and fossil fuel at the power plant gate,ηbiofuel and ηfossil are the electrical efficiencies of the power plant run on biofuel and fossilfuel, respectively, cCO2 is the cost of carbon certificates and fCO2,fossil is the emission factorof the fossil fuel in question.

Figure 4.23 shows the additional cost per MWhel and the GHG mitigation cost for ADM,ADP and HTC in relation to the feedstock price. In addition to HTC-3.00, the resultsfrom simulation cases HTC-3.90 with anaerobic digestion of the waste water and HTC-4.00 with a high ash waste, namely MOW, are also displayed. The carbon certificate pricecCO2 is assumed to be 15 €/t. The additional costs per MWhel are lowest for ADM-3.1,the biomethane production with a high yield, over the entire range of feedstock costs.ADM-3.0 is more sensitive to the feedstock cost than HTC. When the feedstock costs arenegative, the Δc of ADM-3.0 is lower than those of HTC-3.00 and HTC-3.90. When thefeedstock cost exceeds 20 €/t, however, ADM-3.0 is the most expensive of these threescenarios.

Of all scenarios, HTC-4.00 is the most expensive over the entire range of feedstock costs,showing that the feedstock quality can have a strong influence on the Δc. Although notanalyzed here, the same would probably be true for ADM. ADP lies in between ADM-3.0and ADM-3.1 over the entire range of feedstock costs. The relative economic performanceof ADP solid biofuel pellets and ADM therefore depends predominantly on the biomethaneyield of the feedstock. Compared to HTC, ADM performs better in most conditions withthe assumed carbon price of 15 €/tCO2, due to the replacement of a more expensive fossilfuel.

The GHG mitigation costs are, however, lower for HTC than for ADM in most conditions.The GHG mitigation costs of ADM-3.0 are by far the highest over the entire range offeedstock costs. ADM-3.1 performs better than HTC at feedstock costs below −10 €/t,but is more expensive than even the worst performing HTC scenario, HTC-4.00, withfeedstock costs above 10 €/t. ADP-3.0 has the lowest GHG mitigation cost of all theconsidered cases over the entire range of feedstock costs.

The relative merits of solid biofuels and biomethane do not solely depend on the costcriteria analyzed here. Biomethane is a more versatile fuel, and, unlike solid biofuels, canbe used as a transport fuel, in gas boilers for heating applications and in reciprocatingengines for small CHP plants. With increasing amounts of fluctuating wind and solar

158


Figure 4.23: Surplus cost per MWhel when using biofuels instead of fossil fuels (left) andGHG mitigation costs (right), both in relation to the feedstock cost. In the left diagram,carbon certificate costs of 15 €/tCO2 are included.

energy in the electricity generation mix, there may be a shift from coal-fired to moreflexible gas fired power stations. Biomethane might therefore fit better into the futureenergy system than solid biofuel pellets. To answer these wider questions, regional andglobal energy system models are required. Such analysis lies outside the scope of thiswork.

159

5 Bioenergy with carbon capture (BECCS)

As indicated in section 2.3.3, biomass gasification with carbon capture (one form of BECCS)can be pursued by either adapting a typical fluidized-bed biomass gasification system tothe requirements of carbon capture. Or by adapting the biomass to the requirements ofentrained-flow coal gasification systems using torrefaction or HTC. In this chapter, therelative merits of both pathways are analyzed and compared.

First, syngas production with pre-combustion carbon capture is investigated over a rangeof conversion pathways. Fresh wood chips are used as the feedstock to the overall con-version chain in all cases. The analyzed pathways include pretreatment with several ofthe biomass upgrading processes described in chapter 4, as well as some exploration ofthe operating parameters and flowsheet design modifications associated with the syngasproduction.

Second, complete IGCC configurations are analyzed using selected syngas productioncases. Cost estimates are performed, and the electricity production costs of the variousconversion pathways are compared.

Finally, in section 5.3, the biomass-fired IGCC plants are compared to other power planttechnologies with and without carbon capture, including the use of biomethane in a CCGTpower plant.

5.1 Syngas production

The syngas production processes under consideration comprise gasification, air separation,drying, sizing and pressurizing of the feedstock, syngas conditioning, and CO2 separationand compression to 110 bar. The thermal energy released during syngas cooling is utilizedto raise steam at various pressure levels as a byproduct. The final syngas, consistingmostly of hydrogen, could either be used to fuel a gas turbine in an IGCC, or be upgradedfurther as a chemical feedstock or transport fuel.

Table 5.1 gives an overview of the analyzed simulation cases. The column labelled “pre-treatment” indicates the simulation case from which biofuel composition, conversion ef-ficiency and auxiliary energy demand are obtained for the calculation of the conversionchain efficiency. The assumptions for the respective simulation cases are explained in moredetail in sections 5.1.1 and 5.1.2.

For simulation cases with entrained flow gasification, the upgraded biofuel is assumedto be produced in several decentral plants and gasified in a central plant with an inputcapacity of 2250 MWHHV. The fluidized bed gasification is assumed to have a wood chipsinput capacity of 480 MWHHV.

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Chapter 5 Bioenergy with carbon capture (BECCS)

Table 5.1: Simulation cases for syngas production with CCS.

name pretreatment commentsentrained flow gasification

EF-HTC-1 HTC-1.00EF-HTC-102 HTC-1.02 strongly carbonized biocoalEF-TOR-1 TOR-1.0EF-wood-1 none feeding system for direct EF-gasification of milled wood

fluidized bed gasificationFB-wood-1 noneFB-wood-2 none gasification modelled at chemical equilibriumFB-wood-3 none autothermal reformingFB-WP-1 WP-1.0 wood pellets as feedstock

5.1.1 Design and simulation model of syngas production with entrained flowgasifier

A typical plant arrangement for coal gasification with CCS is employed for the entrainedflow gasification of torrefied wood and HTC biocoal. The flowsheet is shown in Figure 5.1.Data for the flow streams for case EF-HTC-1 is presented in Table C.3. Plant design andoperating parameters are largely based on [300, case B.1].

33

31

2

1

88

89

7987

3094

95

201 202

8384

85

85

7832

27

3436 37 35

4 5 6 11

47

28

48

12 13

1415

1716

7173

74

18 72 20 21 69

6768

26

97 205207

208

206

203204

102

98

19

70

57

58

109

108

103

104

105

7

89

10

3959 51

53555654

5260 44 5043 49

4546

coal mill+ drier

lockhopper

coldbox

exhaustgas

cleangas

biocoal

scrubbereffluent

MP

MP

MP MP MP MP

MP

LP

LP

VPMP

HP HP

LP

LP

LP

air

N2

N2

O toClaus plant

2

O2

O2

CO2

N2

gasifier

slag

AGR

Claus

condensate

condensate

shift shift

W1

W5

W4W6W8

W3

W7

W11

W10

solid fuel

steam

syngas

acid gas

flue gas

electricity

CO2

liquid water air, O , N2 2

K1

K28

K29

K20K21

K30

K26 K27K22

K19

K4

K2

K5

K3

K8

K6 K7

K11 K12 K32 K13 K14 K18 K25

K31

K16

K17

K9

A

A

Figure 5.1: Flowsheet of the syngas production process with entrained flow gasifier EF-HTC-1.

The biocoal is milled, dried to 5% water content and pressurized with lock hoppers. Theoxygen-blown gasifier is operated at 1550°C and 39 bar.1 The oxygen is provided by acryogenic air separation unit (ASU). Nitrogen from the ASU is used for pressurizing thelock hoppers and as a drying agent. Some medium pressure steam is generated fromcooling the membrane walls of the gasifier (60). The raw gas leaving the gasifier is cooledto below the slag solidifying temperature (900°C) by mixing it with a colder raw gas recyclestream (K2). It is then further cooled to 310°C and cleaned in a wet scrubber (K6). Thewaste heat is used to generate high pressure (HP) and medium pressure (MP) steam at

1The maximum operating pressure is limited by the performance of the lock hoppers, 30–40 bar areconsidered feasible [249, page 197].

162


137 and 44.5 bar, respectively. The gas leaves the scrubber at 164°C and is reheatedto 275°C before entering the two stage sour shift reactor (K9, K12). More HP and MPsteam is generated from cooling the gas between the two shift stages. The shifted syngasis cooled to 40°C for the acid gas removal (AGR), where sulphur components and CO2 areremoved by physical absorption with the solvent Selexol (K25). The H2S rich acid gas isconverted to elemental sulphur in a Claus plant unit. The separated CO2 is compressedfor transport and sequestration. Thermal energy from syngas cooling after the secondshift stage is used to raise MP, LP (7.2 bar) and VP (4.4 bar) steam and for feedwaterpreheating (K18). Auxiliary energy demand comprises electricity for compressors, fuelmilling, ASU and AGR unit, steam as reactants for the gasifier and water gas shift, andthermal energy for the AGR, ASU and drier, provided in the form of LP and MP steam.

If the plant processes only low sulphur biofuels, hot gas desulphurization may be feasibleand more efficient. However, a design with cold gas desulphurization is chosen here tomaintain the flexibility for utilizing higher sulphur fuels. This would include fossil coaland biocoal from feedstocks with a high sulphur content.

The flowsheet design remains unchanged for case EF-TOR-1, fuelled on torrefied wood,and case EF-HTC-102, fuelled on a more severely carbonized HTC biocoal. With a higherheating value of 29.1 MJ/kg (d.b.), the biocoal in EF-HTC-102 comes close to bituminouscoal. In case EF-wood-1, pulverized wood is gasified without pretreatment. This impliesthat the required feeding technology has been developed and, in this regard, representsa hypothetical case. Based on [251], it is assumed that the feeding system comprises ascrew feeder and a piston compressor and that the high reactivity of wood particles allowsparticles of 1.0 mm to be gasified. The flowsheet is identical to Figure 5.1, except that thewood is first dried to 10% water content, and then milled, because the energy requirementfor milling wood is strongly dependant on its moisture content. Low pressure saturatedsteam at 105°C is used to fuel the drier.

The simulation results are discussed in sections 5.1.3 to 5.1.6.

5.1.2 Design and simulation model of syngas production with fluidized bedgasifier

For fluidized bed gasification with carbon capture, a flowsheet design and Aspen Plussimulation from [346] were refined and adjusted to the general assumptions used in thiswork. The flowsheet is shown in Figure 5.2. Data for the flow streams is presented inTable C.4.

The wood chips are dried to 10% water content and fed to the pressurized gasifier (K5),where they are gasified in a sand bed with a mixture of steam and oxygen. The steammass flow is adjusted so that the raw gas has a steam/CH4 ratio of 2.0, as required by thedownstream methane steam reformer. The gasifier is operated at 900°C and 33.3 bar.2Heat for the wood drier (K1) is supplied by low pressure steam. Size reduction of the woodchips before gasification is not required. Particles are assumed to be removed from theraw gas stream by a cyclone and/or filter,3 before tar and methane are converted to CO

2The operating pressure is determined by the fuel inlet pressure of the IGCC gas turbine (see section 5.2.2)plus pressure losses in the syngas conditioning section.

3This functionality is included in the gasifier block K5 in the simulation.

163


289

287

201

304

308

234305

293 299

306, 236

276 277

238 233292

208

288 290

211 213

254280

257259

260253

214 218 221 224223 226

271

227

261 265262266264

269263

270

295

272

297

273 281

222 225

230

232231

241

294

296

220

219

255, 256

215

216 217

207

274275

252

244

247

250

drier

feedingsystem

coldbox

exhaustgas

wood

air

air N2

N2

O2

CO2

N2

gasifier+particlesremoval

bed 1

bed 2

sulfur

AGR

condensate

water

reformer shift shift

W1

W3W2W12

W9

W5

W4

W7

solid fuel

steam

syngas

flue gas

electricity

CO2liquid water

air, O , N2 2

K5 K6K23

K32

K27 K28

K22 K21K1

K33

K7 K9 K12 K15K13

K16

K24

K25

K26

K17

K10

K29 K8

K11 K14

K19

K20

hot exhaustgas

G

235condensate

300

303

302

301242

205 298

209 239ash

air

sand

dolomite

W11W10

K31

K30

cleangas

LP

MP

MP

HP HP LP

LP

LP

LP

HPHP

Figure 5.2: Flowsheet of the syngas production process with fluidized bed gasifier FB-wood-1.

and H2 in a catalytic steam reforming process (K6). In order to prevent coke formationin the steam reformer, the tar content must be limited to approximately 2 g/m3 (STP)[299]. This is assumed to be achieved by the in-bed use of dolomite in the gasifier, whichis reported to reduce tar formation by up to 95% [258, page 17]. A reduction of tar from12 g/m3 to 2–3 g/m3 (STP) has been achieved in laboratory-scale experiments for thesteam/oxygen gasification of pine wood chips [299].

The sulphur resistance of steam reforming catalysts increases with temperature [266], sothe reformer is operated at 950°C,4 even though this requires a further heating of thesyngas after gasification. Thermal energy for the endothermic steam reforming reactionis provided by burning part of the clean gas. The combustion gas (242) leaves the steamreformer at 970°C and is considered a byproduct of the process. It can, for example, beused to raise steam in the heat recovery steam generator of an IGCC.

Desulphurization to below 0.1 ppmv H2S is required to avoid catalyst poisoning in the lowtemperature water gas shift reactor (K14) [249, page 350]. Adsorption to zinc oxide is wellsuited for almost complete H2S removal given that the inlet H2S content of the gas streamis low [249, page 343]. The analyzed plant design contains a desulphurization unit (K29)with two ZnO beds operated at 480°C and 400°C, as suggested in [348]. Regeneration ofthe sorbent is possible but not considered in this work.

The two stage clean shift (K11, K14) is followed by CO2 removal with the solvent Rectisol(K19). As in the entrained-flow case, a cryogenic ASU (K22) provides oxygen for thegasifier and nitrogen for the feed pressurization system. Thermal energy from syngascooling is used to raise steam at 110 bar (HP) and 6.2 bar (LP).5

Some electricity is generated by the syngas expander (K25), which expands the syngasfuelling the reformer furnace from 17.6 bar to 1.5 bar.

The employed syngas conditioning concept is not proven technology. A similar processdesign to the one suggested here has been analyzed with the help of two small pilot plants inthe CHRISGAS project [266], aimed at producing a gas for the synthesis of transport fuels.

4Typical operating temperatures for steam reforming lie in the range of 800–950°C [347].5The different pressure levels for the steam generation result from the IGCC designs described in

section 5.2.

164


The researchers paid particular attention to the methane steam reforming step. Due to thehigh temperature required for the steam reforming, it is best situated directly downstreamfrom the gasifier, However, deactivation of the steam reforming catalyst by contaminantsmust be avoided. Sulphur poisoning can be prevented by increasing the temperatures inthe steam reformer. However the ash melting temperature must not be exceeded if ashparticles have not been sufficiently well removed upstream [349, page 26]. To this end, acandle filter for particulate removal operating at above 800°C was tested [266, page 12].Methane steam reforming at 900°C reached conversion rates close to chemical equilibriumfor syngas from clean birch wood. Syngas from straw and miscanthus, however, led todeactivation of the catalyst due to its higher sulphur content [266, pages 6, 10]. In-bedsulphur removal in the gasifier with sorbents may be an option [266, page 10]. BesidesH2S, carbon precipitation was another cause for deactivation of the reformer catalysts[349]. These findings illustrate some of the potential difficulties that may be encounteredwhen integrating a methane steam reforming step into the syngas conditioning.

Based on the results from the CHRISGAS project and other data for the individual clean-ing steps retrieved from literature, the plant in Figure 5.2 would appear to be a workabledesign. However, it is highly likely that a considerable R&D effort would be required toreduce all contaminants to the requisite limits and enable a sustained continuous operationfor all process steps on an industrial scale.

The model for the fluidized bed gasification is calibrated using measured data from the IGTgasifier (see section 3.3.2.11 and Table A.21). As discussed in the following section, theresulting methane content from this model is very high. However, the syngas compositiondepends on many factors such as residence time, gasifier design, and chemical reactionswith the bed material. It is therefore questionable whether the syngas composition ofthe commercial scale gasifier with an input capacity of 95 t/h can reliably be predictedbased on limited data from a 0.3–4 t/h [297, 298] pilot plant. Given that the methanecontent has an important impact on the performance of the overall process, an additionalsimulation case FB-wood-2 is modelled with chemical equilibrium in the gasifier.

Case FB-wood-3 employs autothermal steam reforming. Instead of heating the reformer bycombusting clean gas externally, oxygen is mixed into the raw gas in the reaction chamber.The partial combustion taking place provides the thermal energy for the steam reforming.Flowsheet and simulation data for FB-wood-3 are given in Figure C.1 and Table C.5.

Case FB-WP-1 uses wood pellets as the fuel instead of raw wood.

5.1.3 Gasifier efficiency and syngas composition

Key simulation results for the gasification of wood and biocoal for the various simulationcases are given in Table 5.2. Case FB-wood-3 and FB-WP-1 are not listed, because theirgasifier performance is identical to FB-wood-1.

The oxygen demand of entrained flow gasification is 2.5 to 4 times as high as that ofFB-wood-1. To reach the high gasification temperature of 1550°C, a stronger oxidation ofthe fuel is required. This results in a lower cold gas efficiency (CGE) for entrained flowgasification when used with the same fuel. The cold gas efficiency of EF-wood-1 is only69%, compared to 81% for FB-wood-1 and FB-wood-2. For the exergetic efficiency of the

165


Table 5.2: Gasification agents, gasifier outlet composition, HHV of the raw gas, cold gasefficiency and exergetic efficiency of the gasifiers.

EF EF EF EF FB FB

HTC-1 HTC-102 TOR-1 wood-1 wood-1 wood-2

oxygen [kg/kgfeed] 0.741 0.798 0.597 0.534 0.215 0.355

steam [kg/kgfeed] 0.060 0.060 0.060 0.060 0.369 0.060

CO2 [mol%] 5.52% 3.74% 8.97% 12.02% 26.21% 15.34%

CO [mol%] 52.11% 56.66% 41.43% 35.88% 8.28% 32.03%

H2 [mol%] 25.34% 26.40% 23.18% 22.00% 10.72% 29.58%

H2O [mol%] 10.45% 6.77% 19.53% 28.69% 35.83% 18.18%

CH4 [mol%] 0.00% 0.00% 0.00% 0.00% 18.09% 3.70%

N2 [mol%] 6.56% 6.40% 6.88% 1.38% 0.82% 1.12%

H2S [mol%] 0.02% 0.02% 0.02% 0.02% 0.00% 0.00%

tar [mol%] 0.00% 0.00% 0.00% 0.00% 0.06% 0.05%

CGE [–] 78.0% 80.0% 73.4% 69.3% 81.3% 80.9%

εgasifier [–] 83.2% 84.1% 80.8% 79.8% 83.4% 81.1%

HHVraw gas (w.b.) [MJ/kg] 10.35 11.20 8.55 7.70 9.14 10.20

gasifier, however, the difference is not as big. This is due to the higher temperature andassociated physical exergy of the entrained flow gasifier raw gas, which is accounted for aspart of the product by the exergetic efficiency, but is neglected by the definition for thecold gas efficiency.

The feedstock quality has a significant influence on the cold gas efficiency of entrained flowgasification, which ranges from 69% for wood to 80% for strongly carbonized biocoal. Theexergetic efficiency follows the same tendency and ranges from 80% to 84%. The lower thecalorific value of the feedstock, the higher is the CO2 and H2O content in the syngas, andthe lower the syngas higher heating value. The methane content resulting from entrainedflow gasification is below 0.01% in all simulation cases.

The CH4 yield of the fluidized bed gasification is strongly dependent of the simulationmodel used for the gasification. FB-wood-1 results in a very high methane content of18%, in contrast to FB-wood-2, where the CH4 content is only 3.5%. The oxygen demandof the gasifier is 65% higher for FB-wood-2, because the endothermic steam reformingreaction largely takes place in the gasifier itself rather than in the downstream catalyticreactor. Due to the low methane content in FB-wood-2, very little steam is needed forthe catalytic steam reformer. The steam fed to the gasifier is therefore reduced to theminimum value of 0.06 kg/kgfeed.

The methane content of 18% in case FB-wood-1 seems very high when compared to valuesreported in literature. Moreover, in-bed use of dolomite for tar reduction is assumed inthe simulation, but the gasification model is calibrated using data from a gasifier withoutdolomite. Since dolomite was found to decrease the methane content [299], the methanecontent in FB-wood-1 is quite likely overestimated. The methane content in FB-wood-2,on the other hand, is unrealistically low, because at such low operating temperatures,chemical equilibrium will not be reached. Regarding the methane content in the syngas,FB-wood-1 and FB-wood-2 can therefore be considered upper and lower bounds, with a

166


realistic value being somewhere in-between.

Figure 5.3 shows the CH4 content for various pressures and steam/fuel ratios as a functionof the gasification temperature. CH4 formation decreases with temperature and increaseswith pressure. In order to reduce the need for downstream catalytic steam reforming, thegasifier operating conditions should be chosen with the aim of minimizing the CH4 yield.However, the temperature is limited by the ash melting temperature of the feedstock,since ash sintering must be avoided. The pressure is largely determined by the subsequentuse of the syngas and cannot be reduced significantly. Increasing the steam/fuel ratiois technically possible but will decrease the gasifier efficiency. Injecting more steam intothe gasifier than needed to match the H2O/CH4 ratio required by the catalytic reformerhas a negative impact on the exergetic efficiency of the overall syngas production [350].The influence of the gasifier design and in-bed additives on CH4 formation should also beinvestigated further.

0%

5%

10%

15%

20%

25%

30%

35%

700 750 800 850 900 950 1000T [°C]

mol

% C

H4

33.3 bar, rS=0.369, Tapp

20.0 bar, rS=0.369, Tapp

34 bar, rS=0.34, IGT

20-22 bar, rS=0.6-0.7, IGT

20 bar, rS=0.60, Tapp

33.3 bar, rS=0.369, equil

Figure 5.3: Methane content of the syngas as a function of the gasification temperature.Experimental data from the IGT gasifier [297, 298] and simulation data calculated withapproach temperatures (Tapp) (see Table A.21) and under chemical equilibrium (equil).The steam ratio (rS) is given in [kgH2O/kgfeed].


Figure 5.4 shows the energy input required to produce 1 MJHHV of clean, decarbonizedsyngas. The diagram shows the feedstock input, the biomass conversion loss at the pre-treatment plant (i.e. the difference between raw biomass input and upgraded biofueloutput) and the electricity consumed in the syngas production process and in the pre-treatment process. A breakdown of the electricity consumption is shown in Figure 5.5. Inmost simulation cases, the thermal energy recovered from gas cooling is higher than theinternal consumption, resulting in a net output of steam and/or hot exhaust gas as poten-tially useful byproducts. Detailed data for the energy balance of each simulation case is

167


given in Table C.1. Cold gas efficiencies (CGE) for the production of clean, decarbonizedsyngas are summarized in Table 5.3 for the syngas production process and for the overallconversion chain including the pretreatment of the biomass by conventional pelletizing,torrefaction or HTC.

-0.5

0.0

0.5

1.0

1.5

2.0

2.5

EF-HTC-1

EF-HTC-10

2

EF-TOR-1

EF-woo

d-1

FB-woo

d-1

FB-woo

d-2

FB-woo

d-3

FB-WP-1

ene

rgy

inpu

t [M

J/M

J syn

gas ] thermal energy

electricity (pretreatment)

electricity (syngas process)

biomass conversion loss, pretreatment

gasifier fuel

Figure 5.4: Energy input for the production of 1 MJ of clean, de-carbonized syngas(HHV).

Although the cold gas efficiency of the gasifier itself is highest in cases FB-wood-1 andFB-WP-1, the CGE of the overall syngas production process is the lowest in these twocases at 65%. The CGE is highest for EF-HTC-102, where 79% of the gasifier feed energyis recovered in the form of clean syngas. The low syngas yield in FB-wood-1 and FB-WP-1is due to the catalytic steam reforming step, where 31% of the clean gas is combusted toprovide thermal energy for the reformer. In FB-wood-2 this is only 6%, resulting in aCGE more than 10 percentage points higher. Autothermal reforming in case FB-wood-3leads to an even higher increase in CGE, resulting in 77.5%.

Table 5.3: Cold gas efficiencies (CGE) for the production of clean, decarbonized syngas,for the syngas production process and for the overall conversion chain including thepretreatment of the biomass.

EF EF EF EF FB FB FB FB

HTC-1 HTC-102 TOR-1 wood-1 wood-1 wood-2 wood-3 WP-1

syngas production 77.1% 79.0% 72.4% 68.2% 64.9% 76.6% 77.5% 64.6%


If the conversion losses in the biomass upgrading plant are taken into account, the conver-sion pathways with pretreatment suffer a disproportionately high penalty. The CGE forthe conversion chain from raw biomass to clean syngas is only 54% for FB-WP-1, which

168


-0.020.000.020.040.060.080.100.120.140.160.180.20

EF-HTC-1

EF-HTC-10

2

EF-TOR-1

EF-woo

d-1

FB-woo

d-1

FB-woo

d-2

FB-woo

d-3

FB-WP-1

elec

trici

ty c

onsu

mpt

ion

[MJ e

l /MJ

syng

as] pretreatment

drier fan

milling & pressurizing

ASU & oxygen compression

air, flue gas and recycle compressors

AGR

CO2 compressing and drying

SG-Expander

Figure 5.5: Breakdown of the electricity consumption for the syngas production cases.

combines the less efficient gasification process with pretreatment. However, it is also theprocess with the largest amount of thermal energy as a byproduct. Overall efficiencies,taking into account byproducts and auxiliary energy, are discussed in section 5.1.6, basedon exergy. Conversion chain CGE for pathways with entrained flow gasification and pre-treatment ranges from 58% for EF-TOR-1 to 62% for EF-HTC-1. The advantages ofa more efficient syngas production process are outweighed by the conversion losses frombiomass upgrading.

The electricity consumption of the syngas production ranges from 0.07–0.14 MJel/MJsyngas.It is higher in the entrained flow gasification cases, due to the higher oxygen demand re-quiring compression work and because of feedstock milling and pressurizing.6 Again, pre-treatment of the biomass provokes a relatively high penalty. In some cases, the upgradingplant is responsible for more than 30% of the total electricity consumption.


For BECCS systems, the carbon capture rate might be equally or even more importantthan the energetic and exergetic efficiencies. Table 5.4 shows the carbon balance for therespective simulation cases.

While cases with direct gasification of wood result in 82%–86% of the feedstock carbonbeing captured, carbon losses during the pretreatment reduce the overall capture rate to66–69%. For the fluidized bed gasification, 2% of the feedstock carbon remains in the ash,while for the entrained flow gasification, 0.3% ends up in the gasifier slag. This carbonmay also remain in a stable form for a long time, but is not accounted for as “captured”in this analysis.

6In EF-wood-1, electricity demand for milling and pressurizing is lower than in the other entrained flowgasification cases. This is due to the larger particle size after milling and pressurization by pistonfeeders which exhibit a lower inert gas consumption.

169


Table 5.4: Carbon balance for the syngas production cases including pretreatment, basedon the carbon contained in the raw biomass.


HTC-1 HTC-102 TOR-1 wood-1 wood-1 wood-2 wood-3 WP-1

captured 69.4% 66.4% 68.5% 85.8% 82.4% 83.4% 85.8% 69.1%

ash / slag 0.2% 0.2% 0.2% 0.3% 2.0% 2.0% 2.0% 1.7%

not captured during

syngas production

12.1% 11.7% 11.5% 13.9% 15.6% 14.6% 12.2% 13.0%

released at upgrading

plant

18.3% 21.6% 19.8% — — — — 16.2%

The capture rate at the syngas production plant is 82%–86%. Although 90% of the CO2 isassumed to be separated in the absorption unit, some carbon slips through as CO becauseof incomplete conversion in the shift reactor. In the cases with fluidized bed gasification,7–34% of the feedstock carbon is present as methane and would bypass the capture withoutdownstream catalytic steam reforming. This would, of course, severely impair the capturerate.


Results from the exergy analysis are displayed in Figure 5.6. Details are given in Table C.2.Definitions for the exergetic efficiencies are explained in section 3.5.2.

The total exergetic efficiency of the syngas production process according to Equation A.1is 63–66% for raw wood, 69% for wood pellets and torrefied wood, and 71–72% for HTC-biocoal. For raw wood, the process with entrained flow gasification performs worse thanthe process with fluidized bed gasification.

In all simulation cases, the gasifier and reformer are the largest sources of exergy destruc-tion, accounting for 47–71% of the total exergy destruction and loss ED + EL , or 13–22%of the total exergy input to the syngas production plant. In FB-wood-1, about a thirdof this exergy destruction is caused by the catalytic steam reformer, which operates at arelatively low exergetic efficiency of 55%. The efficiency of the fluidized bed gasificationcases is therefore strongly dependent on the methane concentration in the raw gas. Forthis reason, the exergetic efficiency of case FB-wood-2 is 2.8 percentage points higher thanthat of FB-wood-1. The exergy destruction in the steam reformer is reduced by almost halfwhen autothermal reforming is employed instead of an externally heated reformer. Thisleads to case FB-wood-3 being equal to FB-wood-2 in relation to overall process efficiency.

The drier and drier exhaust gas together account for 14–15% of ED + EL in the caseswhere raw wood is used as the feedstock. This explains the higher efficiencies in the caseswhere upgraded biofuels are used as a gasifier feedstock, as the conversion losses of dryingare then born by the pretreatment plant.

The exergy analysis confirms the high efficiency penalty from the various biomass pre-treatment processes, which is discussed in section 5.1.4 in respect to CGE and electricityconsumption. The analyzed conversion pathways employing biomass upgrading result in

170


0

20

40

60

80

100

120

140

160

180

EF-

HTC

-1

EF-

HTC

-

EF-

TOR

-1

EF-

woo

d-1

FB-w

ood-

1

FB-w

ood-

2

FB-w

ood-

3

FB-W

P-1

exer

gy [M

W]

ED + EL otherED AGR + CO2-compED steam and HW generationED wood drier + EL drier exhaustED ASUED steam reformer + shift reactorED gasifier (incl. gas quench)captured CO2steam + reformer exhaust gassyngaselectricity, air, water in upgradingspent biomass in upgradinggasification agents + airelectricitygasifier fuel

56.3%(71.0%)

54.8%(71.9.%)

55.6%(68.5%)

62.9% 63.6% 66.4% 66.4%

68.7%(56.9%)

Figure 5.6: Exergy flows for the syngas production cases, normalized to 100 MWex gasifierfuel. The left column for each case shows the fuels to the syngas production plant (solidcolour) and the exergy destruction and losses of the upgrading plant (hatched). Theright column shows the products (solid colour, red borders) and the exergy destructionand losses of the syngas production plant (hatched). The numbers indicate the exergeticefficiency of the overall conversion chain (bold) and, if differing from the former, of thesyngas production process (in brackets).

40–53% of the overall conversion chain ED+EL occurring during the pretreatment process.The resulting conversion chain exergetic efficiencies range from 55–57% for the pathwaysincluding pretreatment, being highest for wood pelletizing with fluidized bed gasification.More severe HTC results in a slightly higher efficiency for the syngas production due to thehigher quality of the biocoal, but the additional biomass spent in the upgrading processresults in a worse performance for the overall conversion chain.

Given the relatively similar efficiencies and the uncertainties regarding the performanceof some process steps, no clear preference can be given to any of the pretreatment andgasification technologies from a thermodynamic point of view. However, one can concludethat, in respect to efficiency, direct gasification of the raw wood seems to be better thanemploying a separate biomass pretreatment process. The economic merits of pretreatmentare analyzed in the next section for IGCC power plants.

If the biomass supply logistics allow, the integration of biomass pretreatment and syn-gas production may well offer the potential to improve the conversion chain efficiency.However, such arrangements are not pursued in this work.

171


Development of industrial-scale fluidized bed gasifiers for BECCS processes should aim tokeep the methane content of the raw gas low by optimizing the gasifier design and operatingconditions and/or by using in-bed additives. Externally heated steam reforming imposesa high efficiency penalty on the overall process, since a high share of the produced cleansyngas is consumed internally. Burning raw gas rather than clean gas is likely to improvethe efficiency but will lead to significant carbon slippage unless the resulting CO2 can becaptured from the combustion gas. Autothermal catalytic reforming may provide a moreefficient alternative to externally heated reforming.

5.2 IGCC power plants

IGCC plants with CCS were simulated for selected syngas production cases, supplementingthe processes described in the previous section with a combined cycle gas turbine burningthe syngas.

The simulation cases considered comprise entrained flow gasification of HTC biocoal andtorrefied wood, and fluidized bed gasification of wood chips and wood pellets. Whenupgraded biofuels are used as the fuel, two scenarios are distinguished for the cost calcu-lations. In the first, the biofuel is produced in a medium-scale upgrading plant from localSR wood. In the second, the upgrading plant is located overseas and utilizes inexpensiveforest residues. The costs of the respective biofuel production processes are as presentedin chapter 4. For the IGCC fuelled by raw wood chips, only local SR wood is consideredas a feedstock, given that the overseas transport of raw wood is not economically viable[78, page 77]. The fluidized bed gasification of wood chips without carbon capture andthe entrained flow gasification of bituminous coal with carbon capture are also analyzedas reference cases. An overview of the analyzed cases is provided in Table 5.5.

Table 5.5: Simulation cases for IGCC plants.

name pretreatment wood supply commentsBECCS with entrained flow gasification

EF-IGCC-HTC-1 HTC-1.00 m-SREF-IGCC-TOR-1 TOR-1.0 m-SR, l-FR

BECCS with fluidized bed gasificationFB-IGCC-wood-1 none SRFB-WP-1 WP-1.00 m-SR, l-FR

reference casesFB-IGCC-wood-0 none SR without carbon captureEF-IGCC-coal — — bituminous coal

5.2.1 Design and simulation model of an IGCC with entrained flow gasifier

The flowsheet for the IGCC plant with entrained flow gasification is shown in Figure 5.7.Contrary to the syngas production process described in the previous section, the ASU isoperated at an elevated pressure and partially integrated with the gas turbine system.

172


About 30% of the required air is taken from the gas turbine air compressor (131). Someelectricity is recovered by expanding the preheated clean syngas (203) from 27 bar to 17.6bar (K87) before it enters the gas turbine combustion chamber (K34). Nitrogen from theASU is saturated with water vapour (K42) and fed to the gas turbine for NOx control,resulting in a fuel mixture with 50 mol% H2.

The gas turbine combustor exit temperature is 1350°C. Details of the gas turbine modelare provided in section 3.3.2.14.

The gas turbine exhaust gas at 605°C is used to generate steam for a three pressurereheat steam cycle. Details on the operating parameters assumed for the steam cycle aregiven in Table 3.8. Steam generated in the gasification island is fed into the drums at therespective pressure levels, while medium pressure steam for the gasifier and shift reactorand low-pressure steam for the ASU, AGR and drying of the gasifier feed is extracted. Allclean condensate is returned to the deaerator (K66). Treatment of scrubber effluent andcondensate from the syngas conditioning is not considered in the simulation. Steam at 4bar for the deaerator is provided by one of the syngas coolers (K14).

Due to material constraints, the syngas coolers are utilized for evaporation only, but notfor superheating. This results in a high level of integration between the HRSG and theother process units. The HRSG consists mostly of superheaters and economizers, whilethe evaporation takes place in the syngas coolers. The only steam generation takingplace in the HRSG equates to 9% of the high pressure steam. The MP and LP steamconsumption of the process units exceeds their generation, and the net demand is coveredby extractions from the steam turbine (K62). Steam production and consumption issummarized in Table C.10.

Flow stream data for EF-IGCC-HTC-1 is given in section C.2.1.

5.2.2 Design and simulation model of an IGCC with fluidized bed gasifier

The flowsheet for the IGCC with fluidized bed gasification is shown in Figure 5.8. Theclean decarbonized syngas (302) is mixed with steam and nitrogen (K72) to reduce the H2content to 50 mol%, and burned in the gas turbine combustion chamber. Compared tothe entrained flow IGCC, a smaller gas turbine with a lower combustor exit temperatureof 1230°C is employed. The hot exhaust gas from the steam reformer (242) is mixed withthe gas turbine exhaust gas. This raises the gas inlet temperature of the heat recoverysteam generator (HRSG) to 616°C, enabling a relatively high live steam temperature of596°C. A two pressure reheat HRSG produces HP and LP steam, while MP steam for thegasifier and shift reactor, as well as 1.2 bar steam for the drier, is obtained from steamturbine extractions. Similar to the EF-IGCC, the bulk of the steam is produced in thesyngas cooling section. Table C.14 gives an overview of the steam flows.

Since experience concerning the air-side integration between the ASU and the gas tur-bine system in existing IGCC plants is mostly limited to larger gas turbines, no air-sideintegration is included in this plant design.

Flow stream data for FB-IGCC-wood-1 is given in section C.2.1

In the IGCC without carbon capture, FB-IGCC-wood-0, the high methane content of thesyngas does not represent a problem and steam reforming is not required. The gasifier is

173


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174


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175


therefore operated with a lower steam to feed ratio of 0.1. The raw gas is then desulphur-ized and fed directly to the gas turbine combustion chamber. Flowsheet and simulationdata for FB-IGCC-wood-0 are given in section C.2.3.

5.2.3 Energy and carbon balance

Key results for the IGCC cases and their respective conversion chains are given in Table 5.6.

Table 5.6: Efficiencies and carbon capture rates for the IGCC cases and their respectiveconversion chains.

EF EF FB FB FB EF

HTC-1 TOR-1 wood-1 WP-1 wood-0 coal

IGCC

gross efficiency, HHV [–] 44.9% 44.0% 38.6% 41.4% 44.3% 44.6%

net efficiency, HHV [–] 34.2% 32.9% 30.6% 34.2% 41.5% 33.8%

carbon capture rate [–] 84.9% 85.3% 82.3% 82.5% 0.0% 84.5%

conversion chain

net efficiency, HHV [–] 24.3% 25.3% 30.6% 26.9% 41.5% 33.8%


net CO2 emissions [kg/MWhel] −915 −869 −864 −825 0 144

The gross energetic efficiency according to Equation 3.38 is 44–45% for the entrained flowgasification of coal and upgraded biofuels. It is slightly higher for EF-IGCC-HTC-1 thanfor EF-IGCC-coal because of the higher ash and sulphur content of the bituminous coalwhen compared to the HTC biocoal.

The fluidized bed gasification of wood pellets without carbon capture also results in agross efficiency of 44%, while with carbon capture, it is 3–6 percentage points lower. Thehigher gross efficiency of entrained flow gasification is mostly eroded by its higher electri-city consumption for the ASU and for feedstock milling and pressurizing. Regarding netefficiency, EF-IGCC-HTC-1 and FB-IGCC-WP-1 are on a par. The internal electricityconsumption amounts to 17–25% of the generation for the plants with carbon capture butonly 6% for FB-IGCC-wood-0, where CO2 compression and acid gas removal are not re-quired. Since methane is not a problem in FB-IGCC-wood-0, less steam is introduced intothe gasifier, which reduces the oxygen demand and related ASU electricity consumption.The penalty for carbon capture is 10.9 percentage points, comparing the net efficiency ofFB-IGCC-wood-0 and FB-IGCC-wood-1.

Carbon capture rates for the IGCC are identical to those for the respective syngas pro-duction cases discussed in section 5.1.5. For CCS plants fired by fossil fuels, the specificCO2 emissions per unit of produced electricity are often used as a performance metric.Assuming that the biomass is CO2 neutral, the specific emissions are negative for theBECCS cases, ranging from −825 to −915 kg/MWhel. Care needs to be taken with theinterpretation of this value. As long as it is positive, the lower the better. However,when it becomes negative, it is merely a ratio between the two products,7 namely CO2

7Analogous to the power to heat ratio of a CHP plant.

176


removal from the atmosphere and electricity. A lower value in this case is not necessarilybetter. For example, increasing the efficiency of a BECCS plant would result in a higherelectricity output per unit of spent biomass, thereby resulting in less CO2 captured perunit of electricity. However, producing more electricity with the same amount of biomassand the same amount of CO2 removed from the atmosphere would clearly represent animprovement.

When comparing the efficiency of the EF and FB cases, it should be kept in mind thatthe EF design is based on existing coal-fired IGCC plants (albeit without CCS) whilethe FB design is mostly based on laboratory-scale research. It can therefore be expectedthat, by the time a large-scale plant of the FB design is realized, the EF design will haveexperienced some technical progress as well, which may result in an improved efficiency.Moreover, the EF plant design and operating parameters are based on experience fromcoal gasification. Tuning the process to the properties of HTC-biocoal or torrefied woodmay offer some potential to improve the efficiency. The upgraded biofuels have a higherreactivity than bituminous coal, which may allow for a lower gasification temperature.If only low sulphur biofuel is to be processed, hot gas desulphurization instead of lowtemperature absorption could present an attractive option. However, this would severelylimit the choice of feedstock, which may not be desirable.

Compared to the modelling results developed earlier for the same plant configurations[22], the conversion chain efficiency of EF-IGCC-HTC-1 is somewhat lower, because ofdifferences in the HTC plant model (see section 4.5.12.2). The efficiency of FB-IGCC-wood-1, on the other hand, is higher, due to changes in the gasification model and gasifieroperating conditions.

Table 5.7 shows results from other studies for comparison. Three categories are distin-guished: (1) FB-IGCC with CCS, (2) FB-IGCC without CCS and (3) EF-IGCC withCCS. Published simulation studies on biomass-fired IGCC with carbon capture are relat-ively uncommon, whereas studies on plant designs without carbon capture are somewhatmore usual. Numerous simulation studies on coal-fired IGCC, with and without carboncapture, have been published. The data presented in Table 5.7 is not intended to becomplete, but rather focusses on the studies which most resemble the simulation casesanalyzed in this work in terms of gasifier type and plant capacity. A visual representationof the results is provided in Figure 5.12.

Reported efficiencies within each category cover a wide range, often spanning 10 percentagepoints or more. Compared to other studies of biomass-fired FB-IGCC, both FB-IGCC-wood-1 and FB-IGCC-wood-0 lie in the middle of their respective categories. However,results cannot be directly compared because different gasifier types and feedstocks areemployed. The efficiencies for the EF-IGCC cases also lie in the mid-range of studies witha comparably high carbon capture rate of 85% or more. Meerman et al. [351] reportefficiencies 3–9 percentage points higher for EF-IGCC fuelled with torrefied wood andbituminous coal, but with a low capture rate of 56–60%.

177


Table5.7:

Com

parisonofthe

modelling

resultsfor

biomass

andcoal-fired

IGC

Cfrom

thisw

orkw

iththose

fromother

studies.capacity

efficiency

capturerate

feedstock,description

[MW

el ](H

HV

) 1)[%

C]

water

content

FB

-IGC

Cw

ithC

CS

thisw

ork,FB

-IGC

C-w

ood-1143

30.8%82%

wood,50%

pressurized,oxygen-steamblow

ngasifier

+steam

reformer

Larson

andJin,2006

[352]352

35.8%90%

switchgrass,20%

pressurized,oxygen-blown

gasifier

Rhodes

andK

eith,2005[248]

11025.0%

55%biom

assdual-bed,steam

blown

gasifier+

steamreform

er

Carpentieriet

al.,2005[353]

19131.0%

*80%

wood,15%

atmospheric,air-blow

ngasifier

modelled

underchem

icalequilibrium

Audus

andFreund,2005

[354]24

29.3%*

85%w

oodoxygen-blow

ngasifier

FB

-IGC

Cw

ithou

tC

CS

thisw

ork,FB

-IGC

C-w

ood-0194

41.5%0%

wood,50%

pressurized,oxygen-blown

gasifier

Jinet

al.,2009[355]

44245.0%

0%sw

itchgrass,20%pressurized,oxygen-blow

ngasifier

Rhodes

andK

eith,2005[248]

14934.0%

0%biom

assdual-bed,steam

blown

gasifier

Craig

andM

ann,1996[356]

13239.7%

0%w

ood,38%pressurized,air-blow

ngasifier

EIA

,2010[244]

2027.6%

0%w

ood,25%dual-bed,steam

blown

gasifier

Audus

andFreund,2005

[354]30

36.6%*

0%w

oodair-blow

ngasifier

EF

-IGC

Cw

ithC

CS

thisw

ork,E

F-IG

CC

-coal761

33.8%85%

bituminous

coaldry

feed,oxygen-blown

gasifier

thisw

ork,E

F-IG

CC

-TO

R-1

74032.9%

85%torrefied

wood

dryfeed,oxygen-blow

ngasifier

(Shell)

Meerm

anet

al.,2011[351]

80040.0%

56%bitum

inouscoal

dryfeed,oxygen-blow

ngasifier

(Shell)

Meerm

anet

al.,2011[351]

66839.4%

60%torrefied

wood

dryfeed,oxygen-blow

ngasifier

(Shell)

IEA

,2003[300]

67632.9%

*85%

bituminous

coaldry

feed,oxygen-blown

gasifier(Shell)

Carbo

etal.,2007

[357]400

36.1%*

85%bitum

inouscoal

dryfeed,oxygen-blow

ngasifier

(Shell)

DO

E/N

ET

L,2010

[246]497

31.2%90%

bituminous

coaldry

feed,oxygen-blown

gasifier(Shell)

1)values

marked

with

*are

convertedfrom

LH

V-based

values,assuming

HH

V/L

HV

=1.05

forbitum

inouscoaland

1.09for

wood

with

15%w

atercontent.

178



As pointed out in section 5.2.1, there is quite a high level of integration between the HRSGand the other process units which produce and consume steam at various pressure levels.

In Figure 5.9, the HRSGs for cases EF-IGCC-HTC-1 and FB-IGCC-wood-1 are repres-ented with the help of pinch-analysis-style composite curves. The hot and cold compositecurves combine the enthalpy rate difference and temperature characteristics of all hot orcold streams, respectively [358]. The shape of the composite curves allows one to identifythe pinch points of the process. The area between the hot and cold composite curves isproportional to the rate of exergy destruction in the heat exchangers.

As the HRSG is dominated by preheating and superheating rather than evaporation, theslopes of the hot and cold curves are better matched than in a typical combined cycleHRSG. The potential for efficiency improvement in the HRSG therefore seems ratherlimited.

0 200 400 600 8000

100

200

300

400

500

600

700

delta H [MW]

T [°C

]

0 50 100 150 2000

100

200

300

400

500

600

700

delta H [MW]

T [°C

]

EF-IGCC-HTC-1 FB-IGCC-wood-1

Figure 5.9: Composite curves for the HRSG for EF-IGCC-HTC-1 (left) and FB-IGCC-wood-1 (right).

Results from the exergy analysis of selected IGCC plants are visualized in Figure 5.10.The syngas production is responsible for over 50% of the total exergy destruction and lossED + EL of the IGCC in the cases with carbon capture, and 37% in FB-IGCC-wood-0.The gas turbine system contributes 27–44% to the overall ED + EL of the IGCC, whilethe steam cycle at 6–8% has a minor impact.

The exergetic efficiency of the BECCS-IGCC at 34–39% is higher than that of EF-IGCC-coal at 33%. This is because the captured CO2 is accounted for as a product when itresults in a net removal from the atmosphere, but is considered a waste stream whenit originates from a fossil fuel. Regarding the overall conversion chain, pathways withbiomass upgrading exhibit an efficiency of only 29–31%, and FB-wood-1 turns out tobe the best performing case with CCS. Conversely, FB-wood-0 without carbon captureoutperforms all CCS cases.

179


0

20

40

60

80

100

120

140

160

EF

IGC

C-c

oal

EF

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C-H

TC

-1

FB

IGC

C-w

ood-

0

FB

IGC

C-w

ood-

1

FB

IGC

C-W

P-1

exer

gy [M

W]

EL, other

EL exhaust gases

ED steam cycle

ED gas turbine

ED syngas production

captured CO2

electricity

electricity, air, water in HTC

spent biomass in HTC

gasification agents + air

biomass/biocoal to gasifier

33.1%

28.9%(38.9%)

39.2% 34.5%

31.4%(38.4%)

Figure 5.10: Exergy flows for selected IGCC cases, normalized to 100 MWex of gasifierfuel. The left column for each case shows the fuels to the IGCC (solid colour) and theexergy destruction and losses of the upgrading plant (hatched). The right column showsthe products (solid colour, red border) and the exergy destruction and losses of theIGCC (hatched). The numbers indicate the exergetic efficiency of the overall conversionchain (bold) and, if different from the former, of the IGCC (in brackets).


The overall capital investment (TCI) is almost 2000 M€ for the IGCC with entrainedflow gasification and around 500 M€ for the IGCC with fluidized bed gasification. ForFB-IGCC-wood-0 without carbon capture, it is 370 M€. The specific capital investmentranges from 1903 €/kWel for FB-IGCC-wood-0 through 2569 €/kWel for EF-IGCC-HTC-1 to 3581 €/kWel for FB-IGCC-wood-1. This means that carbon capture increases thespecific investment by about 90% for the IGCC with fluidized bed biomass gasification.The specific investment for the entrained flow gasification of upgraded biofuels is 24–38%lower than that using direct fluidized bed gasification of wood chips. The breakdown ofthe module cost (CBM) in Figure 5.11 reveals that the methane steam reforming, whichcontributes 10% to the total CBM, is partly responsible for the high cost of fluidizedbed gasification with carbon capture. Another reason for the higher specific investmentcost of the IGCCs with fluidized bed gasification compared to those with entrained flowgasification is their lower capacity and resulting less favourable economy-of-scale. This is,

180


for example, quite pronounced for the accessory electric plant8, acid gas removal and airseparation unit. For the air separation unit, the effect is partly compensated by the loweroxygen demand per kWel of the fluidized bed gasifiers. A breakdown of the investmentcosts for all cases is given in Table C.22, and detailed equipment lists for FB-IGCC-wood-1and EF-IGCC-HTC-1 are presented in Table C.18 to Table C.21.

Figure 5.11: Breakdown of the module cost (CBM) per unit of electricity production forthe IGCC cases.

A breakdown of the levelized cost of electricity (COE) is given in Table 5.8. The costsor revenues of purchasing or selling CO2 certificates are not included. The impact ofthe CO2 price on the COE is discussed in detail in the next section. The dominant costitems for the BECCS plants are the feedstock (48–75%) and the carrying charges (20–41%). The BECCS case with the lowest COE is EF-IGCC-TOR-1-l-FR using importedtorrefied pellets, but at 132 €/MWhel is still 67% more expensive than EF-IGCC-coal.A CO2 price of 54 € is required for EF-IGCC-TOR-1-l-FR to become cost-competitivewith EF-IGCC-coal. Compared to the biomass gasification case without carbon captureFB-IGCC-wood-0, the COE of the BECCS cases are 49–136% higher.

Because of the large impact of the fuel cost on the COE, the source of the feedstock is moreimportant than the type of IGCC plant, and FB-IGCC-WP-1-l-FR using imported woodpellets is the second to cheapest of the BECCS cases. If domestic wood from short rotationis used, direct gasification of the raw wood is better than prior upgrading. Although theelectricity production costs without fuel costs are 26–30% lower for the entrained flowgasification than for the fluidized bed gasification of raw wood, this does not justify thecost of upgrading if the same feedstock is used. Thus, only if the upgrading enablescheaper biomass resources to be tapped, is it economically viable. However, entrainedflow gasification is a more proven technology, thus building a large FB-IGCC as analyzedhere would require more lead in time for R&D and implies a greater uncertainty regardingavailability, efficiency and cost.

8The accessory electric plant comprises transformers, switchgear and emergency diesel generator.

181


Table 5.8: Levelized cost of electricity for the IGCC plants.


HTC-1 TOR-1 TOR-1 wood-1 WP-1 WP-1 wood-0

coal m-SR m-SR l-FR SR m-SR l-FR SR

capacity [MWel] 761 770 740 740 143 160 160 194

specific TCI [€/kWel] 2530 2569 2717 2666 3581 3142 3104 1903

fuel cost [€/GJHHV] 2.52 14.51 9.94 6.90 5.70 9.35 6.97 5.70

carrying charges [€/MWhel] 41.29 42.00 44.38 43.54 58.46 51.31 50.67 31.07

labour [€/MWhel] 0.78 0.77 0.80 0.80 3.75 2.32 2.32 2.67

O&M, material [€/MWhel] 4.96 4.66 5.09 5.09 6.87 5.91 5.91 3.56

consumables1) [€/MWhel] 0.20 0.22 0.27 0.27 0.64 0.33 0.33 0.47

feedstock [€/MWhel] 27.56 156.95 111.86 77.67 68.93 101.16 75.39 50.78

CO2 transport

& storage

[€/MWhel] 4.19 4.29 4.43 4.43 4.59 4.11 4.11 0.00

COE, total [€/MWhel] 78.98 208.88 166.82 131.79 143.24 165.14 138.74 88.551) waste water and ash disposal, chemicals, ZnO for desulfurization

Due to the higher upgrading cost of HTC compared to torrefaction, the COE of EF-IGCC-HTC-1-m-SR is 25% higher than that of EF-IGCC-TOR-1-m-SR. However, HTCmay allow lower quality (and cheaper) feedstocks to be deployed.

If raw wood chips are used as the fuel, the availability of wood in the vicinity of the IGCCmay significantly influence the COE. If the average transport distance rises by 100 km incase FB-IGCC-wood-1, the fuel costs increase by 18% and the COE by 9%.

Figure 5.12 shows the specific investment costs and efficiencies of the IGCC plant modelsin this work in comparison to similar designs from other studies. The cost estimates forthe fluidized bed gasification of biomass without CCS vary over a considerable range,namely from 1050 to 6754 €/kWel. This is probably due to different feedstocks, gasifiertypes and the wide range of plant capacities from 20 MWel to over 400 MWel. Even forplants with a comparable capacity of 20–30 MWel, the cost estimates vary considerably.Audus and Freund report specific investment costs of 2024 €/kWel for an IGCC with anair-blown fluidized-bed gasifier [354], while the U.S. Energy Information Administrationexpects specific investment costs of 6754 €/kWel for an IGCC with a steam-blown dual-bed gasifier [244]. The more than three times higher investment cost of the latter systemmay be partly due to the more complex gasifier design with two reactors. Despite lessdiverse plant designs, capacities and feedstock, the cost estimates for coal-fired IGCCwith carbon capture still vary from 1680–4576 €/kWel.9 The wide range of cost estimatesindicates a high degree of uncertainty regarding the likely costs for all the plant typesunder consideration.

In sum, the specific investment costs estimated in this work lie within the ranges reportedin literature for all analyzed plant designs.

9Some studies encompass several gasifier types and plant designs. If so, only the simulation case witha Shell gasifier and the plant design most similar to the design considered in this work is depicted inFigure 5.12.

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5.3 Comparison with other biomass, coal and gas to power technologies

Figure 5.12: Efficiency and specific investment cost for IGCC plants analyzed in this work(filled markers) and for comparable plants from other studies [243, 244, 246, 248, 300,351, 352, 354, 355, 359](open markers). All plants are equipped with CCS unless other-wise stated.

5.3 Comparison with other biomass, coal and gas to powertechnologies

To place the obtained results into a wider perspective, it is useful to compare the analyzedIGCC cases to other types of power plant with and without carbon capture. Because theeconomic assumptions used are not uniform, a direct comparison across different studiesis not compelling. Therefore, assumptions on efficiency and specific investment costsare made based on data from literature for selected reference plants, and the COE arethen calculated with the same economic parameters as used here for the IGCC plants.The reference plants comprise pulverized coal power stations without CCS fuelled onbituminous coal and on torrefied wood pellets and gas-fired combined cycle power plants(CCGT) with and without post combustion carbon capture, fuelled on biomethane andnatural gas. The efficiency is assumed to be 54.1% (HHV) or 60% (LHV) for CCGTplants without CCS fuelled on natural gas or biomethane. With CCS, the efficiencies are45% (HHV) or 50% (LHV). The efficiency of a pulverized coal-fired power station withoutcarbon capture is assumed to be 43% (HHV) or 45% (LHV).10

Using torrefied wood pellets in a pulverized coal-fired power station is followed by anefficiency penalty of 1.8% due to the lower fuel quality (see section 4.6.2). A capture rateof 85% is assumed for the CCGT with CCS. Key results and the assumptions on thespecific CBM for the reference cases are summarized in Table 5.9. Based on the CBM,the calculation of TCI and COE is conducted as described in section 3.4.11 A breakdown

10The efficiencies assumed here are higher than those in section 4.6.2 because they refer to new builtplants, while those in section 4.6.2 refer to existing facilities.

11Offsite cost are assumed to be 2.5% of CBM for all cases, based on EF-IGCC-HTC-1.

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Table 5.9: Key assumptions and results for the reference plants.

PC PC CCGT CCGT CCGT CCGT

TOR-1

l-FR

coal ADM-3.0 ADM-3.1 nat. gas nat. gas

no CCS no CCS no CCS

power plant

capacity [MWel] 761.2 761.2 770.1 770.1 770.1 770.1


efficiency (HHV) [–] 41.2% 43.0% 45.0% 45.0% 45.0% 54.1%

specific CBM [€/kWel] 1100 1100 1000 1000 1000 500

specific TCI [€/kWel] 1668 1608 1756 1560 1379 715

fuel cost [€/GJHHV] 6.90 2.52 37.49 21.25 6.31 6.31

COE [€/MWhel] 93.27 51.92 348.07 208.74 80.48 57.45

conversion chain

efficiency (HHV) [–] 31.9% 43.0% 15.2% 29.3% 45.0% 54.1%


of the COE is presented in Figure 5.13. In Figure 5.14, the COE is shown in relation tothe CO2 price.

The conversion chain efficiency for producing electricity from biomass is higher via a solidupgraded biofuel pathway then via a biomethane pathway. The higher conversion losses ofthe anaerobic digestion overcompensate for the higher efficiency of the CCGT in contrastto the IGCC, unless a very high conversion efficiency is achieved in the anaerobic digestion(case ADM-3.1 ). For BECCS applications, the carbon capture rate of the conversion chainmay be equally or even more important than the efficiency. Klein et al. report resultsfrom a global energy system model which indicate that the carbon capture rate has tosignificantly exceed 60% for BECCS to be useful [18]. The capture rate for biomethaneproduction and combustion in a CCGT is only 15–29%, compared to 68–69% for solidbiofuel production and use in an IGCC, and 82% for direct gasification of raw wood. Dueto its low carbon capture rate, anaerobic digestion as a pretreatment technology for aBECCS process does not appear attractive.

Given that a CO2 transport infrastructure exists, the carbon balance of anaerobic digestioncould be improved by capturing the CO2 formed during anaerobic digestion. Assuminga capture rate of 85% for the biomethane production plant would potentially increasethe overall capture rate to 42–61%. If anaerobic digestion is to be further consideredas a pretreatment for BECCS, the influence of carbon capture on the auxiliary energyconsumption and on the costs of the biomethane plant would need to be investigated. Asdiscussed in section 4.4.5, 27–49% of the feedstock carbon ends up in the digestate at thebiomethane plant. Part of this is converted to humus when applied to the soil. Accountingfor this form of carbon storage may further improve the carbon balance, especially whenconsidering a time horizon of less than 100 years.

For all power plants processing biomass, the fuel represents the biggest cost. This iseven more pronounced for the biomethane cases than for the cases using raw biomassor solid biofuels, due to the high cost of the biomethane production. Under the very

184


Figure 5.13: Breakdown of the COE for various power plant types.

optimistic assumptions regarding the biomethane yield made for case ADM-3.1, the COEof biomethane production followed by combustion in a CCGT with CCS is approximatelythe same as for HTC followed by use of the biocoal in an IGCC with CCS. However, thecarbon capture rate is less than half. In case of the lower (and more realistic) assumptionregarding the biomethane yield in case ADM-3.0, the COE of biomethane with CCGT isat least 65% higher than all other analyzed BECCS cases.

A rising CO2 price leads to an increasing COE for power plants using fossil fuels with orwithout capture and to a decreasing COE for BECCS plants with net negative emissions.The relation between the COE and the CO2 price is shown in Figure 5.14.

The break-even point of coal with and without CCS lies at a carbon price of 46 €/t. Belowthat, unabated fossil fuel use is the cheapest option. At 50–51 €/t, biomass without carboncapture FB-IGCC-wood-0 and the cheapest BECCS case, EF-IGCC-TOR-1-l-FR, breakeven with coal without CCS. At 54 €/t, the COE of EF-IGCC-TOR-1-l-FR falls belowthat of coal with CCS. At 64 €/t, EF-IGCC-TOR-1-l-FR breaks even with the unabatednatural gas CCGT and has from then on the lowest COE of all the considered cases.The next best BECCS scenarios, fluidized bed gasification of imported wood pellets ordomestic short rotation wood chips, require CO2 prices about 6–10 €/t higher than EF-IGCC-TOR-1-l-FR to become competitive with the various gas and coal scenarios.

The results indicate that if the carbon price climbs high enough to prevent unabatedfossil fuel use, then the best case BECCS scenario, entrained flow gasification of importedtorrefied wood pellets, is also not far from economic viability.

The influence of the carbon price on the COE is much stronger for the BECCS scenariosthan for fossil fuel use with CCS. At a CO2 price of 50 €/t, for example, the revenuesgenerated by the negative emissions offset 32% of the electricity production cost, or 41.6

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Figure5.14:

CO

Ein

relationto

thecarbon

price.T

heright

diagramis

anenlargem

entofthe

most

relevantarea

oftheleft

diagram.

186


€/MWhel, for EF-IGCC-TOR-1-l-FR. This means that the revenues from the negativeemissions are about equal to the carrying charges and are therefore one of the determiningfactors for the COE. For EF-IGCC-coal, on the other hand, the CO2 cost only contributes7 €/MWhel to the cost of electricity.

A BECCS plant therefore faces two economic risks: the lack of cheap biomass feedstocksand a low carbon price. An IGCC design with entrained flow gasification, apart fromhaving the lowest COE, offers the advantage of being suitable for both coal and upgradedbiofuels. Thus fuel-switching offers some flexibility to react to changing biomass andcarbon prices. The analyzed design of a fluidized-bed IGCC is less flexible regarding thefuel choice. Due to the lower operating temperature, switching to fossil coal with its lowerreactivity is probably not suited. In addition, the hot gas cleaning requires a clean fuelwith low sulphur, chlorine and other contaminants. This requirement will rule out fossilcoal and probably also biomass feedstocks other than untreated wood.

For entrained flow gasification, waste biomass upgraded with HTC could provide an al-ternative to imported torrefied pellets. Depending on the remuneration rate for wastetreatment, the cost can be similar (see section 4.6).

Due to the lower carbon content of biomethane compared with solid biofuels, the COE ofa CCGT fuelled with biomethane decreases less rapidly with a rising carbon price. Evenat a high CO2 price, the revenues generated by the negative emissions therefore do notoffset the high biomethane cost, and the COE remains the highest of all the technologypathways considered here.

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6 Conclusions and outlook

Hydrothermal carbonization (HTC) converts raw biomass into a lignite-like substance,which is better suited for combustion or gasification and which also has favourable prop-erties for transport and storage. Extensive laboratory-scale research on the HTC reactionhas been presented in the scientific literature, but published analysis concerning the en-ergy balance and economic viability of industrial-scale HTC plants is largely lacking. Asof early 2013, there are no commercial HTC facilities in operation.

This work provides an assessment of the techno-economic performance of industrial-scaleHTC plants in a variety of contexts. To this end, HTC is compared to competing biomassupgrading technologies and to the combustion and gasification of raw biomass. The keymetrics are energetic efficiency, GHG emissions and costs. The upgrading technologiesanalyzed in this work comprise wood pelletizing, torrefaction, biomethane production viaanaerobic digestion (ADM) and the combined production of biogas and a solid biofuel(ADP). Simulation models of the processes are created with the software package AspenPlus, which allows the auxiliary energy demand and waste streams to be quantified, theplant equipment to be sized, and the investment costs and the cost of the biofuel productto be estimated.

Wood from domestic short rotation plantation and forest residues imported from NorthAmerica, source-separated municipal organic waste (MOW), park and gardening waste(PGW) and empty fruit bunches (EFB) from palm oil production are considered as feed-stocks. A simple supply chain model was developed and used to assess the GHG emissionsand costs of biomass procurement and transport.

Exergy-based methods were applied to analyze the HTC process in detail and to identifypotentials for improvement in the flowsheet design and operating conditions.

The first part of this work focusses on the short term replacement of fossil fuel by biofuelsin existing power stations. The second part investigates the potential role of biomassupgrading technologies for bioenergy with carbon capture and storage (BECCS).

6.1 Comparison of different biomass upgrading technologies

The energetic efficiency (based on HHV) of solid biofuel production from wood is quitesimilar for wood pelletizing (82%), torrefaction (79%) and HTC (78%). For solid biofuelproduction from grass and waste biomass the efficiency is somewhat lower, 71% for ADPand 68–75% for HTC, due to the higher ash and moisture content of these feedstocks.The energetic efficiency of biomethane production is only 41% when hard-to-digest grassis used, due to the unconverted organic material which leaves the process as digestate.Biomethane yields can vary significantly depending on the time of harvest. Optimistic

189

Chapter 6 Conclusions and outlook

assumptions on the yield can lead to efficiencies up to 75%, which is in the same range asthe solid biofuel production processes.

Production costs are similar for wood pellets and torrefied pellets, ranging from 8.5 to 11.4€/GJ when produced from domestic short rotation coppice. Producing the pellets fromforest residues in North America and shipping them to Europe leads to a cost reductionof approximately 2 €/GJ. HTC biocoal pellets are 50–85% more expensive than woodpellets or torrefied pellets produced from the same feedstock, due to the higher invest-ment and operating costs of the more complex upgrading process. However, while simplepelletization and torrefaction are mostly suited to relatively dry feedstocks, HTC is alsowell suited to wet waste biomass. Treatment of park and gardening waste may attractremunerations of around 20 €/t, in which case HTC biocoal pellets can be produced at8.4–13.8 €/GJ. The lower end of this range is cost-competitive with wood pellets. UsingMOW as the feedstock, with a remuneration rate of 50 €/t, can lead to negative biocoalproduction costs. However, the biocoal pellets from MOW are of low quality due to theirhigh ash content of over 20%. The revenues received for waste treatment are equal to orhigher than the revenues from selling the biocoal. In this context, HTC is more a wastetreatment technology than a biofuel production process. For a comprehensive analysis ofits market potential, HTC needs to be compared to other waste treatment technologiessuch as composting and incineration. Such an assessment lies outside the scope of thiswork.

Investment and labour costs for an HTC plant are strongly dependent on the plant size. Aminimum processing capacity of around 100 kt/a is required to make HTC biocoal pelletsfrom PGW cost competitive with wood pellets. This is equivalent to the annual PGWof approximately two million people, which means that locations with sufficient feedstocksupply are most likely to be found in densely populated urban areas. While specific in-vestment and labour costs decrease with increasing plant capacity, the feedstock transportcosts increase due to the larger required catchment area and related transportation dis-tances. Analysis shows that the optimal processing capacity for HTC from PGW is around150–250 kt/a, or 30–50 MWHHV.

Some biomass processing industries create large amounts of waste biomass, which is oftenavailable year-round at one site. Due to legal requirements, most waste in Germany isutilized or adequately treated, but dumping of organic residues is common practice inmany countries. One example of an unutilized bioenergy potential is the waste frompalm oil production. A typical palm oil mill delivers 20–90 kt/a of empty fruit bunches(EFB), which represents an attractive feedstock for HTC. Including sea transport fromMalaysia, HTC biocoal pellets could be delivered to Europe at 10.11–11.6 €/GJ. Besidesunlocking the potential of unutilized bioenergy, treating the EFB with HTC can avoidlarge amounts of methane emissions if it replaces dumping. If the avoided GHG emissionsare remunerated at 18 €/tCO2, the HTC biocoal pellets can be delivered to Europe at 7€/GJ, similar to the cost of imported wood pellets from North America.

All biofuels produced from grass are expensive because of the high feedstock cost. Pelletsfrom ADP at 20 €/GJ are cheaper than those from HTC, which amount to 25 €/GJ. Theproduction costs of biomethane are largely dependent on the biogas yield of the anaerobicdigestion and range from 20–37 €/GJ using the assumptions in this work. Given thatbiomethane is a higher quality product than solid biofuels, it seems that biomethane

190

6.1 Comparison of different biomass upgrading technologies

production is economically favourable with moderate to high biogas yields, whereas solidbiofuel production could be better suited to feedstocks with a low biogas yield, such asthe harvest from conservation grassland.

The most widely discussed use of torrefied wood pellets and HTC biocoal pellets is co-combustion in coal-fired power stations. Substituting fossil fuels by biofuels in existingpower stations offers power plant operators some flexibility to react to rising CO2 pricesand may provide a short term option to increase the use of bioenergy. HTC biocoal pro-duced from MOW in plants with a processing capacity of 93 kt/a or more can break evenwith bituminous coal, without accounting for avoided carbon certificates. No other biofuelscenario is cost competitive with fossil coal unless CO2 prices increase markedly, or the re-spective technologies receive support under feed-in tariff or renewable certificate schemes.In Germany, the co-firing of biomass with coal is not covered under the Renewable EnergySources Act.

When substituting bituminous coal in an existing power station, the GHG mitigation costfor a medium-scale HTC plant (48 MW) processing green waste amounts to 77 €/tCO2,eq.This is similar to a medium-scale wood pelletizing plant supplied with short rotationcoppice. Importing pellets produced from forest residues in large-scale pelletizing plantssituated in North America results in 53 €/tCO2,eq.

The GHG mitigation cost for HTC from EFB amounts to 28–34 €/tCO2,eq, whereat theGHG avoidance from substituting EFB dumping and subsequent methane release is largerthan the GHG avoidance from displacing fossil coal combustion at the power plant. Thus,any treatment method replacing EFB dumping has a high GHG reduction potential, andalternatives such as composting, anaerobic digestion or even using the EFB for mulchingthe plantation should be analyzed and compared to HTC.

Replacing natural gas in a CCGT power plant by biomethane from grass leads to veryhigh GHG mitigation costs of 398–917 €/tCO2,eq. The high costs are partly due to theexpensive feedstock. However, the production of solid biofuels from grass, ADP and HTC,lead to significantly lower GHG mitigation costs of 170 and 261 €/tCO2,eq, respectively.This is partly because the GHG reduction is greater when substituting coal than whendisplacing natural gas. The GHG emissions for using solid biofuels in a power station,including the supply chain emissions, amount to 5–22% of those from using fossil coal.For biomethane, they amount to 35–41% of those from using natural gas. However, theGHG emissions from anaerobic digestion reported in literature vary widely, and a moredetailed analysis, including different scenarios for methane losses and the utilization of thedigestate, need to be conducted for a comprehensive assessment.

One advantage of the upgraded biofuels is that they facilitate the large-scale use of bio-mass in existing power stations with no or little modification required at the power plant.However, the investment required to build the upgrading plants is considerable. For woodpelletizing, it equates to 290–830 €/kWel and for torrefaction to 480–1130 €/kWel. Thesecosts are similar to adapting an existing coal-fired power station to enable the co-firingof raw biomass, which has been estimated at 170–1000 €/kWel [11]. HTC at 2200–5900€/kWel and ADM at 2940–6000 €/kWel are both in the same range as the specific in-vestment required for building a dedicated wood-fired power station, which is reported as1360–6750 €/kWel [244, 245, 247, 342].

191


Based on the analysis conducted in this work, no clear preference can be given to any of theupgrading technologies in respect to production costs or energetic efficiencies. However,the influence of the biofuel quality on the power plant operation has not been investigatedin detail. HTC biocoal is the most “coal-like” of the solid biofuels, with the highestcalorific value and probably the best properties regarding storage and feed preparationfor pulverized coal burners. The properties of torrefied pellets lie between those of HTCbiocoal and conventional wood pellets. These considerations may be decisive for theselection of the best suited upgrading process.

The exergetic efficiency of the conversion chain from raw biomass to electricity ranges from20% to 27% for conversion pathways with upgraded biofuels and combustion in a pulverizedcoal-fired power station. These figures are in the same range as for the utilization of rawbiomass in a dedicated biomass-fired CHP plant. The higher efficiency of the power plantis counteracted by the relatively high conversion losses in the upgrading processes. ForADM, the exergetic efficiency of the conversion chain when burning the biomethane in aCCGT ranges from 16–31%, depending on the methane yield of the anaerobic digestion.

A detailed exergy analysis of the upgrading processes reveals that the hydrothermal pro-cesses suffer relatively large exergy losses because of organic compounds contained inresidue streams. The digestate of ADM accounts for 19–52% of the feedstock exergy, andthe dissolved compounds in HTC waste water account for 9%. Drying of the feedstockbiomass or the biofuel, and combustion to provide the thermal energy for drying, repres-ents the biggest source of exergy destruction in all of the solid biofuel production processesconsidered. While wood pelletization and torrefaction require drying of the feedstock bio-mass before the chemical conversion, in HTC the drying takes place after the chemicalreaction. This offers two advantages. Firstly, oxygen is removed from the feedstock beforedrying, thus the amount of water to be evaporated per MJ of biofuel is reduced. Secondly,the structure of the biomass is altered, thereby allowing mechanical dewatering to a greatextent. The exergy destruction from drying is thus about halved when compared to woodpelletizing and torrefaction. On the downside, the exergy destruction within the remainingplant components, including heat recovery, is higher in HTC due to the more complex pro-cess. When wood is used as a feedstock, the overall exergetic efficiency is therefore slightlylower for HTC. For very wet feedstocks such as waste biomass with a water content of70%, however, pretreatment of the biomass with HTC improves the overall conversionefficiency compared to the combustion of the raw biomass, in cases where the combustionfacility does not use flue gas condensation.

6.2 Detailed analysis of the HTC plant design

The HTC base case flowsheet analyzed in this work derives from existing plant designsfor the hydrothermal treatment of peat, of which several pilot-scale facilities were in op-eration until the 1960s. Unit operations comprise the pressurization and preheating ofthe biomass slurry, the hydrothermal reaction, depressurization, cooling and mechanicaldewatering of the biocoal slurry, and drying and pelletizing of the biocoal. Heat recoveryfrom the biocoal slurry and heating of the feed slurry is realized with a counterflow ar-rangement whereby the pressure of the product slurry is stage-wise reduced and the flashsteam condenses in the feed slurry. Organic compounds in the waste water are oxidized

192

6.2 Detailed analysis of the HTC plant design

with an aerobic treatment step. Seven modified flowsheet designs, featuring alternativepreheating schemes, a different arrangement for the biocoal dewatering and alternativedrying technologies, were simulated and analyzed with respect to energetic efficiency andeconomic performance. Additionally, 16 parameter studies were conducted to assess theinfluence of key operating parameters such as the reaction conditions and dry mattercontent at various stages of the process.

The analysis shows conclusively that an efficient heat recovery scheme is essential forreasons of energetic efficiency and biocoal production costs. An arrangement using heatexchangers is potentially more efficient than the direct heat transfer employed in the basedesign. Problems with clogging and fouling of the biomass and biocoal slurry may preventthe use of heat exchangers, but further R&D into the clogging and fouling characteristics ofthe slurries and specialist heat exchanger designs may reveal potentials for improvement.

The investment costs are strongly dependent on the reactor volume, and accordingly on theresidence time of the HTC reaction. “Over-carbonization”, meaning carbonizing more thanis required to reach the desired hydrophobicity and breakdown of the fibrous structure,should therefore be avoided. A higher reaction temperature allows for shorter residencetimes for a given level of carbonization. The results of the parameter studies indicate thatincreasing the operating temperature in order to shorten the residence time reduces thebiocoal production cost. However, more experimental data on carbonization is requiredto confirm this result and to identify the best operating conditions.

The water content in the biomass slurry has a strong influence on the biocoal cost, sinceadditional water flowing through the process means that most plant equipment items mustbe larger and therefore more expensive. The minimum required water to biomass ratiodepends on the feedstock pressurization system employed.

Steam leaving the reactor with the gaseous byproducts significantly impacts the energeticefficiency, because the thermal energy required for evaporation must be compensated forby additional preheating of the biomass slurry. In order to minimize evaporation in thereactor, the ingress of inert gas or air into the reactor must be avoided. This rules outbiomass pressurization systems such as lock hoppers and puts further constraints on thedifficult selection of a suitable pressurization system. Other potential measures to decreasethe evaporation in the reactor are an operating pressure well above saturation pressureand reactor designs where the exiting gases are cooled by direct heat transfer with theincoming biomass.

A high degree of mechanical dewatering of the biocoal slurry decreases the size, investmentcost and energy demand of the drier. Whether or not this also improves the efficiency ofthe overall process depends on the entire plant design. In some configurations there is asurplus of low pressure steam, thus the energy saved in the drier is discharged anyway.

The most significant measures to improve the efficiency are the use of heat exchangersinstead of direct heat transfer, and a design where the biocoal is dried in superheatedsteam. Both arrangements allow the entire steam demand for preheating the biomass tobe covered by internal heat recovery. This leads to energetic efficiencies of 78–79% (HHV),compared to 72.6% for the base case. In most other analyzed cases, 8–14% of the producedbiocoal is burned to produce steam for preheating the biomass.

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The biocoal production costs of the investigated design alternatives and operating condi-tions are in the range of ±20% of the base case.

Depending on the feedstock, HTC produces 2.5–6.5 kg waste water per kg biocoal. Notonly do the dissolved organic compounds represent a significant energy loss, but the associ-ated waste water treatment will most likely add significant costs. Because the compositionof the waste water has not yet been fully identified, a high uncertainty remains about therequired treatment steps and costs. A treatment cost of 5 €/m3, which derives from theestimates presented earlier, translates to 1.1–1.2 €/GJ biocoal, or 9–13% of the overallbiofuel production costs. The high electricity demand for the aerobic treatment is alsoresponsible for around half of the total electricity demand of the HTC plant.

Other than aerobic treatment, anaerobic digestion allows part of the energy in the wastewater to be recovered in the form of methane. The biogas can be utilized to raise steamfor biomass preheating. Experiments with the anaerobic digestion of HTC waste water in-dicate good degradability, although a subsequent aerobic treatment step will probably stillbe required [220, 338]. Adding an anaerobic treatment step to the waste water treatmentincreases the energetic efficiency of the HTC process from 72.6% to 78.5%, and decreasesthe biocoal production cost by 6%.

Integrating an HTC process with a rankine-cycle CHP plant allows steam for the biomasspreheating to be taken from a steam turbine extraction, and enables low pressure steamrecovered from the HTC process to be utilized for district heat supply. Compared tostandalone CHP and HTC plants, the energetic efficiency is roughly the same, but thebiocoal production costs are 25–26% lower. Moreover, the complex heat recovery schemein the HTC process with stagewise pressurization and preheating of the biomass can beomitted, which should improve operability. However, locations with an adequate supplyof waste biomass and the required district heat infrastructure will be limited.

6.3 Bioenergy with carbon capture and storage (BECCS)

Bioenergy with carbon capture and storage (BECCS) is a concept whereby biomass iscombusted in a facility with carbon capture. The net result is a carbon negative processwhich, in effect, removes carbon dioxide from the atmosphere.

The analysis of BECCS plants presented in this work focusses on IGCC with pre-combustioncarbon capture. For IGCC fired on bituminous coal, the preferred technology is pressur-ized oxygen-blown entrained flow (EF) gasification, which produces a tar-free syngas witha very low methane content. For biomass, however, entrained flow gasification is not feas-ible, because its fibrous structure impedes the pneumatic transport and pressurization offine particles. Since torrefaction and HTC destroy the fibrous structure, they may facilit-ate the use of entrained flow gasification for biomass. The relative merits of this conversionpathway are explored by comparing thr energetic efficiency and economic performance ofEF gasification to the fluidized bed (FB) gasification of raw wood, based on process sim-ulation models. The chosen capacities are 140 MWel for the FB-IGCC design and 740MWel for the EF-IGCC design.

One major downside for employing fluidized bed gasification in combination with CCS isthat 7–34% of the feedstock carbon is converted to methane in the gasifier. To prevent this

194

6.3 Bioenergy with carbon capture and storage (BECCS)

carbon from slipping through the water gas shift reactor and thereby escaping the capture,a catalytic steam reformer is required downstream of the gasifier. This steam reformer isresponsible for a high exergy destruction and adds significantly to the investment costs.Up to 31% of the produced clean gas has to be burned in order to provide thermal energyfor the endothermic steam reforming reaction. Development of fluidized bed gasifiersfor BECCS should therefore seek to keep the methane content of the raw gas low byoptimizing the gasifier design and operating conditions and/or by using in-bed additives.Autothermal reforming, whereby oxygen is blown into the syngas stream, may provide amore efficient alternative to externally heated reforming.

While the efficiency of the EF-IGCC is 3.5 percentage points higher than that of theFB-IGCC, the overall conversion chain efficiency from raw biomass to electricity is lower.Conversion losses during torrefaction or HTC pretreatment result in an overall electricalefficiency of 24–25% (HHV), compared to 30.6% for the FB-IGCC fired on raw wood.Since one purpose of BECCS is to “remove” CO2 from the atmosphere, the carbon capturerate may be as important as the electrical efficiency. Carbon losses during the biomassupgrading have a significant impact on the capture rate. As a result, the capture rates ofconversion pathways via biomass upgrading are limited to 66–69%, compared to 82–86%for the direct gasification of wood.

The dominant cost factor for the IGCC plants is the feedstock cost (48–75%). The lowerspecific investment cost of the EF-IGCC cannot compensate for the cost of biomass up-grading. Thus, if domestic short rotation wood is used in both cases, using the raw woodin an FB-IGCC is cheaper. The EF-IGCC fired on imported torrefied pellets, however,leads to the lowest cost of electricity (COE). This indicates that biomass upgrading is onlyworthwhile if it enables cheaper biomass resources to be tapped.

Another important aspect is that entrained flow gasification is a more proven technologythan fluidized bed gasification, especially in combination with CCS. Therefore, buildingcommercial-scale FB-IGCC would require a greater lead time for further R&D.

The analyzed IGCC cases are compared to selected reference processes with and withoutcarbon capture. Assumptions on efficiency and specific investment costs for the referenceplants, comprising pulverized coal power stations without CCS fuelled on bituminous coaland on torrefied wood pellets, and CCGT with and without post combustion carboncapture, fuelled on biomethane and natural gas, are based on data from literature.

The conversion chain efficiency for producing electricity from biomass is higher via solidupgraded biofuel then via biomethane. The higher conversion losses from anaerobic di-gestion overcompensate for the higher efficiency of the CCGT as opposed to the IGCC,unless a very high methane yield is achieved for anaerobic digestion. Moreover, the carboncapture rate for biomethane production and combustion of the biomethane in a CCGTis only 15–29%. The low carbon capture rate makes anaerobic digestion unattractive asa pretreatment technology for BECCS. Given that a CO2 transport infrastructure existsat the biomethane production site, the carbon balance of anaerobic digestion could beimproved by capturing the CO2 removed from the biogas.

With the assumptions used in this work, a CO2 price of 46 €/tCO2,eq is required to makeCCS worthwhile for bituminous coal. A CO2 price of 51 €/tCO2,eq means the COE of thecheapest BECCS case, namely an EF-IGCC fired on imported torrefied wood pellets, falls

195


below that of the equivalent plant fired on bituminous coal. For CO2 prices above 60–70€/tCO2,eq, BECCS plants fired on domestic raw wood or imported biofuel pellets providethe lowest COE of all considered cases. They are cheaper than plants fired on bituminouscoal or natural gas, and also cheaper than biomass-fired IGCC without CCS. Althoughthere is a high uncertainty regarding the absolute cost, these results indicate that if thecarbon price rises sufficiently high to prevent unabated fossil fuel use, then the BECCSconcept also enters the realm of economic viability.

The influence of the carbon price on the COE is much stronger for the BECCS scenariosthan for fossil fuel combustion with CCS. An IGCC design with entrained flow gasificationoffers the advantage of being suitable for both fossil coal and biofuels, thus offering theflexibility to fuel-switch in reaction to changing feedstock and carbon prices.

6.4 Outlook

HTC appears most suitable for waste biomass as a feedstock, because its economic viabilityis dependent on the revenues for waste treatment. Since no experimental data was availablefor PGW and MOW, the carbonization model used in this work was calibrated using datafor wood. Carbonization experiments with different types of waste would help to build amore exact model for feedstock-specific mass and energy yields and for the biocoal quality.Other waste streams not considered here, including sewage sludge, may prove to be viablefeedstocks with large potentials and should be included in any future analysis. HTC needsto be compared to alternative treatment technologies for the respective types of waste inorder to determine its competitiveness.

The treatment of the HTC waste water, which is likely to have a significant impact on theeconomic performance of HTC, needs to be investigated in greater detail.

The cost of transporting the feedstock biomass can strongly influence the overall biocoalproduction cost, especially if wet waste biomass is transported over long distances. De-pending on the supply situation, these transport costs can amount to over 30% of thebiocoal costs from an HTC plant processing 100 kt/a of PGW. Ship transport of biofuelsfrom overseas locations to Europe can present a significant cost item as well. Given theimportance of transportion costs, supply chains deserve to be analyzed in greater detail.This includes location-specific studies of the feedstock availability and a comparison of dif-ferent modes of transport. This is especially relevant when assessing the economic viabilityof a concrete project.

Combining HTC with a CHP plant or with biogas production from the waste water bothseem promising options. Since HTC appears well suited for waste streams from biomassprocessing industries, such as EFB from palm oil production, integration of HTC withthe respective primary processes may offer potentials for lowering the cost of biocoal pro-duction and warrants further analysis. Different options to combine HTC with anaerobicdigestion seem especially interesting. A design similar to the combined biogas and presscake pellet production discussed in this work, where press fluid from the feedstock is con-verted to biogas and the hard-to-digest press cake is converted to biocoal, is one possibleroute. Waste streams from biorefineries may provide a potential source of HTC feedstocksin the future, and may also offer extensive possibilities for process integration.

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6.4 Outlook

HTC and torrefaction both make biomass accessible for entrained flow gasification. Besideselectricity generation with CCS, a much discussed application of entrained flow gasificationis the production of liquid transport fuels via Fischer-Tropsch synthesis. While the analysisprovided in this work is limited to biomass-to-electricity pathways, the potential role ofbiomass upgrading technologies for liquid fuels remains an important field of research.

For a more comprehensive comparison of HTC and anaerobic digestion, a full life cycleassessment (LCA) is indicated. Process-related methane emissions and the environmentaleffects of spreading digestion residues on agricultural land may both have a significantimpact, and are worth a more detailed analysis than that provided in this work.

For BECCS concepts employing anaerobic digestion, the potential to improve the carboncapture rate needs to be analyzed. Besides capture and storage of the CO2 releasedduring the digestion process, combination with a power-to-gas process [360] seems to bean interesting option. Power-to-gas is a concept that provides storage for fluctuating windand solar energy. In the process, the CO2 from the digestion is converted to additionalbiomethane by methanation with H2 produced at times when electricity is in surplus. Asa result, the biomethane yield could be increased and the carbon loss from the digestioncould be decreased.

It is assumed in this work that the solid biofuels replace fossil coal in a pulverized coal-fired power station, and biomethane replaces natural gas in a CCGT plant. Biomethane,however, is much more versatile in its application, being suited also to small co-generationplants and for use as a transport fuel. In a future energy system, where combustionfuels merely supplement wind and solar energy, a gaseous fuel like biomethane may fitbetter into the overall system. While the methodology applied in this work allows one tocompare different conversion technologies which produce a solid biofuel, the assessment ofthe relative merits of solid and gaseous biofuels should be supplemented by a regional orglobal energy system analysis.

197

A Modelling assumptions

A.1 Supply chain of biofuel production

A.1.1 Storage

The annual dry matter loss for SR wood chips Lannual can be calculated based on themonthly loss Lmonth, assuming that the harvest takes place once per year and is consumedby the upgrading plant within t months (here t=12), where an equal amount is taken fromthe storage every month [76]:

Lannual = 1 + t · Lmonth

(1 − 1

1 − (1 − Lmonth)t

)

Monthly dry matter losses for wood chips of 1.8–6.6% are reported in literature [361].2.4% is assumed in this work.

The annual dry matter loss from grass ensilage is based on [272].

Storage costs are estimated based on costs for land, construction, labour and machineryto load and unload the storage, and data from [272] for grass ensilage.

Table A.1: Storage dry matter loss and cost.

SR FR grasstype of storage outside, chips outside, logs silagedry matter loss [% p.a.] 15% 3% 9%cost [€/tDM] 6.62 2.00 31.29

A.1.2 Transport distance and cost

A.1.2.1 Road transport cost

Specific transport costs for the considered vehicles per km and tonne c [€/km/t] are estim-ated based on a fixed price component for charter cost ccharter [€/km], fuel consumptionQfuel [MJ/km], fuel price cdiesel [€/MJ], and freight capacity of the vehicle mfreight [t].

c =ccharter + cdieselQfuel

mfreight

For trucks transporting grass and for agricultural vehicles, the volume capacity is the lim-iting factor, and the maximum freight mass is calculated using the volume capacity of the

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Chapter A Modelling assumptions

vehicle and the density of the freight. Assumptions for charter cost and fuel consumptionare based on [78, p. 68]. Charter costs from [78, p. 68] are multiplied by 2 to includean empty return journey. The assumptions regarding freight capacity and charter costfor the different vehicles are listed in Table A.2. Fuel consumption is assumed to be 18.1MJHHV/km for all vehicles [78, p. 68], and the diesel price is 34.50 €/GJHHV. Deviatingassumptions for truck transport in Malaysia comprise a diesel price of 11.66 €/GJHHV anda 30% lower charter cost. The resulting transport cost for the considered vehicles per kmand tonne are given in Table A.3.

Loading plus unloading the vehicle is assumed to contribute a cost of 0.5 €/m3 [78, p. 68].

Table A.2: Assumptions for road transport costs.

freight charter cost

[€/km]

truck, pellets 29 t 1.70

truck, pellets (Malaysia) 29 t 1.19

truck, raw biomass 23 t, 107 m3 1.70

truck, pulverized coal (dangerous goods transport) 33 m3 2.48

collection vehicle, waste 14 t 1.70

agricultural vehicle 25.4 t, 62.8 m3 1.60

Table A.3: Costs for road transport.

truck, pellets [€/km/t] 0.0802

truck, pellets (Malaysia) [€/km/t] 0.0483

truck, pulverized HTC biocoal [€/km/m3] 0.0941

truck, wood chips [€/km/t] 0.1011

collection vehicle, waste [€/km/t] 0.1660

truck, PGW-70 or grass [€/km/t] 0.1301

truck, PGW-50 [€/km/t] 0.2172

agricultural vehicle, wood chips [€/km/t] 0.1854

agricultural vehicle, grass [€/km/t] 0.2121

loading + unloading [€/m3] 0.5

Calculation of the transport distance from farm gate to upgrading plant for wood andgrass:

mharvest,DM =mfeedstock,DM

1 − Lstorage

Acrop =mharvest,DM

yDM · (1 − Lharvest)

Acatchment =Acrop

f1

dtransport = δ · 23

·√

Acatchment

π+ 2

200


mfeedstock [t/a] feedstock consumed by the upgrading plantmharvest [t/a] feedstock harvestedLharvest [–] mass loss during harvestLstorage [–] annual dry matter mass loss during storageyDM [t/km2/a] annual biomass yield (DM)Acrop [t/a] required cropland area to grow the feedstockAcatchment [t/a] total catchment area for acquiring the feedstock for the upgrading plantf1 [–] percentage of land used for growing the feedstockδ [–] tortuosity factor, the deviation between actual and linear transport distance.

δ = 1.414dtransport [km] average driven distance for transporting feedstock to plant, based on [279, p.

61]

A.1.2.2 Calculation of the transport cost for wood and grass

Transport costs from field to farm gate Cfield are estimated as a function of the croplandarea:

Cfield = 2.05 · Acrop · 1.424

The transport distance of 2.05 km/a/ha is estimated based on a typical field geometry,cost of 1.424 €/km comprise charter and fuel costs for agricultural vehicles as describedin section A.1.2.1.

Transport from farm gate to upgrading plant is conducted by agricultural vehicle (Cagri)or truck (Ctruck), whichever is cheaper. Transport by truck requires reloading from theagricultural vehicle used for the transport on the field to the truck. For longer distances,the cheaper transport by truck more than compensates for the cost for the additionalreloading process (Creload). The specific cost for transport by agricultural vehicle (cagri)and truck (ctruck) in [€/km/t] and for reloading [€/m3] are given in Table A.3. Storageis assumed to take place at the upgrading plant when the biomass is transported to theplant without reloading. When the biomass is reloaded onto trucks, storage and storagelosses are assumed to take place at the collection point at the farm.

Cagri = cagrimharvestdtransport + Cfield

Ctruck = ctruckmfeedstockdtransport + Creload + Cfield

A.1.2.3 Calculation of the transport distance for waste

The model described in Table A.1.2.1 is modified for waste. Since collection and transportto a disposal facility are independent of the treatment technology, no cost and GHG emis-sions are taken into account for the amount of waste processed by an average compostingplant mic, acquired from the inner catchment (ic) area Aic. When the capacity of theupgrading plant is higher than that of an average composting plant, transport cost andGHG emissions are incurred for the additional feedstock moc, which is acquired from the

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outer catchment (oc) area Aoc. It is assumed that the inner catchment area is a denselypopulated urban area, from which 100% of the waste is provided to the plant in question.In the outer catchment area, the population density is assumed to be lower, and only partof the waste can be acquired for the upgrading plant, due to existing contracts for alter-native disposal or utilization pathways. Assumptions for the inner and outer catchmentarea and the availability of waste are listed in Table A.4. Data for the waste generationper person are based on [28, p. 160].

moc = mfeedstock − mic

Aic =mic

fic · pic · y∗

Aoc =moc

foc · poc · y∗

dtransport = δ · 23

·√

Aoc + Aic

π+ 2

p [persons/km2] population densityy∗ [kg/person/a] annual waste generation per personf [–] percentage of waste available for upgrading plant

Table A.4: Assumptions for the calculation of the transport distance of waste to the up-grading plant.

PGW MOWwaste per person y∗ [kgFM/person/a] 48 941)

average capacity of a composting plant mic [ktFM/a] 10 20

inner catchment outer catchmentpopulation density p [persons/km2] 2.177 400waste ratio available for upgrading plant f [–] 100% 20%

1) source-separated municipal organic waste plus park and gardening waste

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A.1.2.4 Shipping cost

Table A.5: Assumptions and results for the shipping cost, based on [78, 362].

Malaysia North America North America

pellets type HTC TOR, HTC WP

distance [km] 16000 7000 7000

dead weight (dwt) [t] 35000 74000 74000

ship data

vessel type Handysize Panamax Panamax

speed [km/h] 27.8 27.8 27.8

charter cost [€/d] 15097 18871 18871

fuel consumption [t/km] 0.0430 0.0620 0.0620

fuel price [€/t] 377 377 377

cargo data

density [kg/m3] 750 750 650

pellets volume [m3] 39667 84000 96923

pellets mass [t] 29750 63000 63000

harbours, canals

loading/unloading speed [m3/h] 300 300 300

loading/unloading cost [€/t] 2.22 2.22 2.22

harbour fees [€/tdwt] 1.11 1.11 1.11

number of harbours [-] 2 2 2

canal fees [€/t] 4.47 0 0

results

charter cost [€/t] 17.76 10.13 11.21

fuel cost [€/t] 8.74 2.60 2.60

harbour and canal fees [€/t] 7.08 2.61 2.61

loading/unloading cost [€/t] 2.22 2.22 2.22

total [€/t] 35.79 17.56 18.63

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A.1.3 Supply chain GHG emissions

Table A.6: GHG emissions from transport.

operation GHG emissions sources and comments

transport

ship [gCO2,eq/t/km] 8.9 overseas shipment, 2005 [275]

truck [gCO2,eq/t/km] 339.8 truck, diesel-mix, Germany, 2005 [275]

pipeline

(biomethane)

[gCO2,eq/t/km] 361.9 calculated using on a natural gas consumption for

compressor stations of 1% per 250 km and a methane loss of

0.3% for transmission and distribution [363]

embedded emissions of electricity

Germany [kgCO2/MWhel] 644 German power mix 2005 [275]

Malaysia [kgCO2/MWhel] 890 Malaysian power mix [364]

fossil fuels

combustionbituminous coal [kgCO2/GJLHV] 95 [365]

pulverized lignite [kgCO2/GJLHV] 99 [365]

natural gas [kgCO2/GJLHV] 56 [365]

embedded emissionsbituminous coal [kgCO2, eq/GJ] 12.91 mix Germany, at power plant gate [275]

pulverized lignite [kgCO2, eq/GJ] 14.67 average of Lausitz and Rhenish lignite [275]

natural gas [kgCO2, eq/GJ] 8.60 mix Germany, at power plant gate [275]

A.2 Process simulation

A.2.1 Modelling assumptions

A.2.1.1 Properties of non-conventional solids

Correlation for the HHV (d.b.) in [MJ/kg], from [286], where c, h, o, n, s, a are the massfractions of carbon, hydrogen, oxygen, nitrogen, sulphur and ash on a dry basis:

HHVdry = 34.91c + 117.83h − 10.34o − 1.51n + 10.05s − 2.11a

Calculation of the LHV (w.b.), from [366, p. 13], where w is the water content on wetbasis:

LHVar = (1 − w)(HHVdry − 21.2h − 0.08(o + n)) − 2.44w

The heat capacity cs,dry [kJ/kg/K] for the dry solids is calculated with the followingcorrelation, where T is the temperature in [K]. The parameters a and b for the varioussolid materials are given in Table A.7.

cs,dry = aT + b

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Table A.7: Parameters for the heat capacity correlation.

biomass biocoal asha [kJ/kg/K2] 0.0039 0.0025 0.00029b [kJ/kg/K] 0.5572 0.5517 0.67312

The parameter b in Table A.7 for biomass is based on the apparent heat capacity of thedry matter for birch wood at full hydration at 25°C [287], and the gradient a for thetemperature dependence is taken from [288, p. 3–17]. The correlation used for ash issuggested by Kirov for ash, slag and clinker [289]. For biocoal, a heat capacity of 1.3kJ/kg at 25°C is assumed based on data for different fossil coals [289, Fig. 3]. Thegradient a is based on the implementation of the Kirov heat capacity correlation in AspenPlus.The physical exergy of a nonconventional solid is calculated using:

eP Hs,dry = (h − h0) − T0(s − s0)

where the enthalpy and entropy differences (h-h0) and (s-s0) are calculated by integrationfrom T0 to T of

dh = cs,dry(T ) dT

ds =cs,dry(T )

TdT

The chemical exergy for the ash-free nonconventional solid is calculated based on theelemental composition and HHV with a correlation for solid fuels with a high oxygencontent [290, p. 105], where C, H, O, and N are the mole contents:

eCHs,daf = LHVdaf

1.044 + 0.016HC − 0.3493O

C

(1 + 0.0531H

C

)+ 0.0493N

C

1 − 0.4124OC

The chemical exergy for the dry nonconventional solid is calculated assuming that theexergy of the ash is zero:

eCHs,dry = (1 − a) · eCH

s,daf + a · eCHash

A.2.1.2 Turbomachinery

The electricity generated by turbines and consumed by compressors and pumps is calcu-lated as follows:

turbine: Wel = Wshaft · ηmechanical · ηgenerator

compressor: Wel = Wshaft/ηmechanical/ηmotor

pump: Wel = Wpump/ηmotor

The shaft work is calculated in Aspen Plus based on the polytropic efficiency. Inter-cooledcompressors in the IGCC plants are assumed to be water-cooled to 30°C between thecompression stages.

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For assumptions on compressor and expander of the gas turbine system, see section A.2.1.12.

Table A.8: Efficiencies of turbomachinery.

sources / commentspolytropic efficiencies

steam turbinesHP section 90%MP section 92%LP section 87%small industrial ST 80% 3–25 MWelcompressorscentrifugalcompressor

ηpolytrop = 0.0098 ln(Vin) + 0.7736 function derived based on datafrom [367, p. 484-485], V in [m3/s]

axial compressor ηpolytrop = 0.0098 ln(Vin) + 0.8336 6 percentage points higher thancentrifugal compressor [306, p. 83]

oxygen compressor ηpolytrop,O2 = ηpolytrop − 0.1mechanical efficiencies

compressor 98%expander 98%

pump efficienciescentrifugal pump ηpump = 0.0497 ln(m) + 0.6097

ηpump,min = 0.60function derived based on datafrom [368], m in [kg/s]

HTC slurry pump see biomass pressurizationelectrical efficiencies

motor ηmotor = 0.0235 ln(Wshaft)+0.7829 function derived based on datafrom [306, p. 93], Wshaft in [kW]

generator, IGCC 98.5% capacity: >30 MWgenerator, CHP 97.0% capacity: 3–25 MW

A.2.1.3 Heat exchangers

Heat exchangers are assumed to be adiabatic. Minimum temperature differences are givenin Table A.9 and pressure losses in Table A.13. The surface area A required for theinvestment cost estimate is calculated with

A =Q

U · ΔTln

where the overall heat transfer coefficients U (see Table A.10) are estimated dependingon the hot / cold fluid based on data from literature [305, 369–371].

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Table A.9: Minimum temperature differences of heat exchangers.

gas / gas (low pressure) [°C] 30gas / gas (high pressure) [°C] 20condensing steam / gas [°C] 10gas / liquid [°C] 10liquid / liquid [°C] 10evaporator temperature difference at pinch point [°C] 10superheater hot gas inlet / steam outlet [°C] 20

Table A.10: Overall heat transfer coefficients for heat exchangers, in [W/m2/K].

biomass slurry water evaporation HP gas LP gasbiomass slurry 1500water 1200 200 100evaporationcondensing steam 1500 2000 2000 200 200HP gas + cond 1) 500 100HP gas 200 200 100 50LP gas 100 20

special heat exchanger types [W/m2/K]air-cooled condenser 800air-cooled cooler 600water-cooled condenser 2500rotary air preheater 20HRSG economizer 50HRSG evaporator 100HRSG superheater 30

1) with partial condensation

A.2.1.4 Heat and pressure losses

Table A.11: Heat losses of process units.

torrefaction reactor [% of torrefaction gas, HHV] 3%steam reforming reactor [% of heat supply] 3%water gas shift reactor [% of heat of reaction] 2%gas turbine combustion chamber [% of fuel, HHV] 1.5%drier [% of heat supply] 5%CHP engine [% of fuel, HHV] 3%HTC piping heat loss, biomass slurry equivalent to 1°C temperature dropHTC piping heat loss, biocoal slurry equivalent to 1°C temperature drop

See also Table A.19 and Table A.20 for heat losses from boilers, furnaces and gasifiers.

207


The heat loss from the HTC reactor was calculated in detail for a preliminary design ofHTC-1.00 and resulted in 14.7 kJ/kg dry biomass, equivalent to 0.08% of the feedstockenergy. Key parameters for the calculations are listed in Table A.12. 14.7 kJ/kg are thenassumed for all HTC cases, except for HTC-3.01 and HTC-3.02, where they are adjustedproportionally to the residence time.

Details on the heat loss calculations of the digesters for anaerobic digestion are given insection B.3.4.

Table A.12: Assumptions for heat loss calculations of the HTC reactor.

reactor design

biomass dry matter [kg/h] 2000

water [kg/h] 11669

reaction temperature [°C] 220

residence time [h] 4

volume, total (liquid phase) [m3] 64.7

number of reactors [–] 2

reactor inside diameter [m] 1.05

reactor height [m] 9.34

parameters for heat loss calculations

wall thickness [mm] 50

insulation thickness [mm] 300

outside temperature [°C] 15

heat transfer coefficient, inside [W/m2/K] 2000

thermal conductivity, wall [W/m/K] 50

thermal conductivity, insulation [W/m/K] 0.1

heat transfer coefficient, outside [W/m2/K] 30

calculated heat loss

heat loss through reactor wall, per reactor [kW] 2.7

factor for piping, connections, instrumentation [kW] 1.5

total heat loss, all reactors [kW] 8.2

heat loss per kg dry biomass [kJ/kg] 14.7

208


Table A.13: Pressure losses of process units.

process unit pressure loss

reactors

HTC reactor [bar] 2.0

torrefaction reactor [bar] 0.12

gas turbine combustion chamber [-] 3%

gas turbine combustion chamber, fuel inlet valve [bar] 2.0

steam reformer [bar] 2.0

water gas shift (per stage) [bar] 2.0

heat exchangers

low pressure gas [–] 0.60% per 100°C temperature change

high pressure gas (> 10 bar) [–] 0.65% per 100°C temperature change

water [–] 2.0% per 100°C temperature change

superheated steam [–] 3.0% per 100°C temperature change

evaporation [–] 5.0%

condensation [–] 0.0%

HRSG, gas side [mbar] 45

HTC slurry mixing preheaters (steam to slurry Δp) [bar] 2.0

gas cleaning equipment

wet scrubber [bar] 1.5

acid gas removal [bar] 3.3

pressurized water scrubbing [bar] 1.0

A.2.1.5 Size reduction

The desired particle size has a large impact on the electricity demand. Milling torrefiedwood to 100 μm requires approximately four times as much electricity as milling to 300 μm[39]. The required particle size for combustion and gasification depends on the technologyused. Generally, combustion and gasification applications with a short residence timerequire a small particle size for the chemical reactions to complete. For a given technology,the required particle size of different feedstocks may vary with the reactivity of the fuel.For entrained flow gasification, bituminous coal must be milled to 90–100 μm [249, p. 130,135]. For wood, which is more reactive than bituminous coal, a particle size of 1 mm wasfound to be sufficient for entrained flow gasification in gasification tests run at the IGCCplant at Buggenum [372, p. 13]. The reactivity of torrefied wood and HTC-biocoal is inbetween that of wood and bituminous coal. Laboratory-scale entrained-flow gasificationof HTC biocoal and lignite indicated a similar reactivity of both fuels at temperaturesabove 1000°C [203]. For fluidized bed gasification of wood, a particle size of up to 70 mmis feasible [28, p. 611]. Short rotation poplar and willow wood is usually cut to 15–45 mmchips during harvesting [28, p. 230]. Therefore, no further reduction is required at thegasification plant.

It is assumed that for HTC and torrefaction, no prior size reduction is required upstreamof the reactor. For wood pellet production, wood chips are assumed to be milled to 6 mm.Due to their brittle nature, it is assumed that HTC biocoal and torrefied wood are brokenup in the pellet mill and do not require milling prior to pelletization.

209


Some experimental results for milling raw wood, torrefied wood and HTC biocoal havebeen published in the scientific literature. An electricity consumption of 135 kJ/kg wasmeasured for milling HTC biocoal to a particle size of 280 μm using a pin mill with athroughput of 100 kg/h [203]. Milling torrefied wood to 200–300 μm requires 100–200kJ/kg with a cutter mill [39]. Due to the brittle nature of torrefied wood, the energyconsumption is 30–44% lower than for milling raw wood on a mass basis. Compared tobituminous coal, it is approximately 30% lower[39].1

While the energy demand measured in laboratory scale milling experiments is useful tocompare different feedstocks, it will deviate from that of large coal mills in a power plant.For example, the electricity demand of large scale coal mills ranges from 36–54 kJ/kg [203],while milling bituminous coal to 180 μm required approximately 360 kJ/kg in experimentswith a cutter mill [39].2

In the analyzed plant designs with entrained flow gasification, biocoal and torrefied woodpellets have to be milled to 100 μm. No data was found in literature for milling such biofuelpellets, but given that the pellets are harder than the unprocessed biocoal or torrefiedwood, the electricity demand for milling is probably higher. 180 kJ/kg is assumed. Forcomparison, Naundorf reports 150 kJ/kg for producing pulverized lignite [373].

The assumptions for milling wood chips and upgraded biofuels are summarized in Table A.14.

Table A.14: Milling electricity demand.

feedstock purpose size reduction to electricitydemand [kJ/kg]

wood chips pelletizing 6 mm 135.6torrefied wood pelletizing in pellet press 0HTC biocoal pelletizing in pellet press 0wood chips fluidized bed gasification none 0biofuel pellets entrained flow gasification 0.1 mm 180.0wood chips entrained flow gasification 1.0 mm 153.4

A.2.1.6 Biomass pressurization

The modelling assumptions and the resulting specific electricity consumption per kg ofdry biomass in the plant models are given in Table A.15.

The pump efficiency of the slurry pump is calculated from vendors data cited in [336]for a piston pump for high-density solids with an electricity consumption of 45 kWel forpumping 10 m3/h biomass slurry to 30 bar.

The specific electricity demand calculated for the plug-forming feeder is based on vendorsdata cited in [336] and [337, p. 91] for a plug-forming piston feeder with an electricityconsumption of 100–130 kWel for feeding wood chips against a pressure of 24 bar.

A maximum dry matter content of 15% is assumed in the simulation model for the slurrypump. This may be a conservative estimate, since such pumps have been successfully

1Electricity consumption is given on an energy basis in [kWel/MWth] in [39], and converted to mass basis.212 kWel/MWth [39], converted to kJ/kg assuming a heating value of 30 MJ/kg.

210


tested for pressurizing peat with a dry matter content of 25%. Vendors state that adry matter content of 30–60% may be feasible [336]. The Swedish Peat Process, on theother hand, was designed for a much lower dry matter content of 7% [200, p. 122]. Themaximum dry matter content strongly depends on the feed material.

The inert gas consumption of lock hoppers and piston feeders for the IGCC plant modelsis specified in Table A.16. The specific inert gas consumption is estimated based on datafrom [251] for the pressurization of 100 μm torrefied wood particles via a lock hopper,and for pressurizing 1000 μm milled wood particles using a piston feeder. It is assumedthat nitrogen at 60 bar is used for the operation of the pressurization systems. For thepressurization of wood chips in the fluidized-bed IGCC, the same inert gas consumptionis assumed as for the piston feeder.

Table A.16: Inert gas consumption for lock hoppers and piston feeders.

particle size total inert gasconsumption

inert gas enteringthe gasifier

[μm] [kgN2/kgfeed] [kgN2/kgfeed]lock hopper 100 0.782 0.135piston feeder ≥1000 0.014 0.005

A.2.1.7 Mechanical dewatering

Assumptions on the mechanical dewatering equipment are given in Table A.17. For thefilter press, it is assumed that ash and organic compounds dissolved in the liquid phaseend up in the solid and liquid products in proportion to the water (thus no additional solidmatter is forced into the liquid phase by the pressing process). The electricity demand ofthe filter press is calculated based on the volume of the displaced water as suggested in[198]: It is assumed that in a first step the biocoal is dewatered to a dry matter contentof 40% with an applied pressure of 30 bar, and in a second step it is dewatered to thefinal dry matter content (60% in the base case) applying a pressure of 100 bar. Therequired work W = pVdisplaced is produced from electricity with an efficiency of 90%.Although the calculation is originally for a piston press, the resulting electricity demandof 3.4 kJ/kgslurry seems plausible. For comparison, the electricity demand for a filter pressprocessing sewage sludge is 3.6–5.5 kJ/kgsludge based on data from [374].

It is assumed that the maximum operating temperature for the filter press is 90°C, thusthe slurry inlet flow is cooled to that temperature. For the high temperature dewateringin HTC-3.50, no cooling is required, but a temperature drop of 2°C is assumed due toheat loss.

211


TableA

.15:E

lectricityconsum

ptionofbiom

asspressurizing.

pressurein

plantm

odel

DM

contentin

plantm

odel

max.

DM

contentpum

peffi

ciencyelectricitydem

andelectricity

demand,

inplant

model

Sourcesand

comm

ents

[bar][kJ/kg

ar ][kJ/kg

feed,d

ry ]slurry

pump

2715%

15%30.5%

71–82piston

pump

forhigh-density

solids[336]

plug-forming

feeder32

30%≥

70%64

213plug-form

ingpiston

feeder[336]

lockhopper

3990%

449based

oninert

gasdem

and[251]

pistonfeeder

34–3990%

10based

oninert

gasdem

and[251]

212


Table A.17: Modelling assumptions for mechanical dewatering.

filter press screw press decanterapplied to biocoal slurry conditioned grass

silageAD fermentation

residuessolid phase DM content 60% 47.30% [234] 30.00% [375]solid flow into liquid phase varies1) 27.7% [234] neglectedelectricity consumption f(Vdisplaced) 14 kWh/tbiomass

[234]1.5 kWh/m3 [375]

1) dissolved organics and ash are proportional to their concentration in the slurry water

A.2.1.8 Thermal drying

Unless otherwise stated in Table A.18, it is assumed that the exhaust leaves the drier at10°C above the dew point Td (saturation temperature) to avoid recondensation.

Modelling assumptions for the atmospheric and pressurized SSD drier are based on [186,293].

The electricity consumption of the drier air fans for the belt drier are calculated with apressure drop of 37 mbar.

Table A.18: Modelling assumptions for driers.

belt drier drum drier atm. SSD press. SSD coal drier(IGCC)

feedstock biomass,biocoal

wood biomass,biocoal

biomass,biocoal

biocoal,fossil coal

inlet water content [–] >30% >30% >30% >30% 10%outlet water content [–] 10% 10% 10% 10% 5%drying medium air hot gas steam steam nitrogenoperating pressure [bar] 1 1 1 3–6.5 1drying medium inlet

temperature[°C] 90 500 240

heat loss [–] 5% 5% 5% 5% 5%exhaust gas outlet

temperature[°C] Td+ 10 110 Tsat Tsat Td+ 10

feedstock outlettemperature

[°C] Td+ 10 80 Tsat Tsat Td+ 10

fluidization steam [kg/kgfeed] 0.5 0.5temperature

difference[°C] 10 1) 20 2) ≥40 2) 10 1)

1) in heat exchanger, between hot water or steam and drying air2) between Tsat (p) of the heating steam and Tsat (p) in the fluidized bed. The pressure of the heatingsteam is calculated to fulfill this condition.

213


A.2.1.9 Boilers and furnaces

Table A.19: Modelling assumption for combustion, based on [314, pp. 30-38] and [28, p.416].

type stoker boiler fluidized bedboiler

furnace incinerator pulverized coalburner

fuel biocoal pellets,wood

biocoal pellets,wood

wood,gas

offgas biocoal pellets,bituminous coal

product saturated orsuperheatedsteam

superheatedsteam

hot gas hot gas superheatedsteam

lambda 1.4 1.2 1.4 1.4 1.15Q-loss 3.0% 2.0% 3.0% 3.0% 0.5%unburned C 3.0% 0.3% 3.0% – 0.3%

A.2.1.10 Gasification

Table A.20: Modelling assumptions for gasification.

entrained flow fluidized bedoperating pressure [bar] 39 33.3operating temperature [°C] 1550 900gasification agent oxygen oxygen + steamunconverted carbon [% C] 0.3% 2%heat loss [% fuel, HHV] 2% 3%membrane walls heat duty [% fuel, HHV] 0.28%pressure loss [bar] 4 4additional pressure loss,cyclone and filter

[bar] 2

The pressure loss in Table A.20 is related to the gasification agents oxygen and steam.

For fluidized-bed gasification, the approach temperatures for the gasification reactions arecalculated based on 3 data sets for the IGT/Renugas gasifier. An Excel tool developed byOgriseck3 is used for the calculation of the approach temperatures based on gas composi-tion, temperature and pressure. The mean value of the approach temperatures resultingfrom the 3 data sets is used for the simulation. Data and resulting approach temperaturesare given in Table A.21.

3Available for download at http://ogriseck.org/download.htm, downloaded 16-Jun-2011.

214


Table A.21: Data from IGT/Renugas gasifier and resulting approach temperatures.

Source [298] [297] [297]

measured measured predicted

temperature [°C] 911 982 968

pressure [bar] 21.9 34.0 20.3

steam [kg/kg feed,ar] 0.690 0.340 0.600

oxygen [kg/kg feed,ar] 0.260 0.255 0.255

gas composition

CO [mol%] 8.43% 15.00% 7.83%

CO2 [mol%] 20.34% 23.90% 17.71%

H2 [mol%] 14.81% 20.80% 15.68%

H2O [mol%] 47.05% 31.80% 50.60%

CH4 [mol%] 8.87% 8.20% 5.73%

other hydrocarbons [mol%] 0.50% 0.30% 0.00%

approach temperatures mean

C + H2O ↔ CO + H2 [°C] -260 -227 -326 -271

CO + H2O ↔ CO2+ H2 [°C] 3 -170 -25 -64

C + CO2 ↔ 2 CO [°C] -216 -213 -277 -235

C + 2 H2 ↔ CH4 [°C] -251 -213 -263 -242

CO + 3 H2 ↔ CH4+ H2O [°C] -257 -225 -303 -262

A.2.1.11 Syngas and biogas cleaning and conditioning

Modelling assumption are given in Table A.23 for the steam reformer, water gas shift, hotgas desulphurization, Claus plant and pressurized water scrubbing, and in Table A.22 forthe acid gas removal units.

Table A.22: Modelling assumptions for acid gas removal units.

H2S + CO2 removal CO2 removal only

data based on unit operated with Selexol

[300, case B1]

unit operated with Rectisol

[376, case B1]

syngas inlet temperature [°C] 38 43

operating temperature [°C] 38 -30

CO2 removal rate [–] 90% 90%

H2S removal rate [–] 100%

H2 slippage into CO2 stream [–] 1.7% 1.7%

CO slippage into CO2 stream [–] 2.5% 2.5%

H2S concentration in acid gas [mol%] 22%

electricity consumption [kJ/kmolraw gas] 3149 1718

LP steam consumption [kJ/kmolraw gas] 4590 1069

CO2 leaving the AGR 30% at 23.3 bar,

25% at 9.2 bar,

45% at 2.9 bar

31% at 10 bar,

34% at 2.7 bar,

34% at 1.4 bar

215


Table A.23: Modelling assumptions for gas cleaning and conditioning equipment.

process unit modelling assumptions

methane steam

reformer

operating temperature: 950°C

minimum H2O/CH4 ratio: 2.0

reactions:

CH4+H2O ↔ CO+3 H2 (equilibrium)

C6H6O+ 6 H2O → 5 CO+9 H2+CO2 (complete conversion)

C7H8+7 H2O → 7 CO+11 H2 (complete conversion)C16H10 +16 H2O → 16 CO+21 H2(complete conversion)

temperature difference between hot combustion gas at outlet and reaction temperature:

30°C

water gas shift,

sour

2 stages at 254° and 242°C (based on [249])

minimum H2O/CO ratio: 2.0

reaction: CO+H2O↔H2+CO2

40% and 15% of CO is set as inert in the first and second stage, respectively, in order to

yield a realistic final CO content of 1.6% [249, p. 350]

water gas shift,

clean

high temperature stage at 344° and low temperature stage at 200°C (based on [249])

minimum H2O/CO ratio: 2.0

reaction: CO+H2O↔H2+CO2

15% of CO is set as inert in each stage, in order to yield a realistic final CO content of

0.3% [249, p. 351]

hot gas

desulphurization

2 ZnO beds: 480°C and 400°C [348]

no energy consumption, sorbent regeneration not considered

Claus plant black box model, characterized by specific syngas and O2 consumption, and steamproductionsyngas consumption: 0.048 kg/kgS, O2 consumption: 0.617 kg/kgS

MP steam production: 0.30 kg/kgS, LP steam production: 0.420 kg/kgS

pressurized water

scrubbing

data based on [232]

inlet pressure: 8 bar

methane loss: 1% (oxidation of the offgas prevents methane emissions)

methane content of the biomethane: 97% d.b.

electricity consumption: 0.14 kWh/m3STP (without compression of the biogas to the

inlet pressure). The electricity consumption has been calibrated to match the overall

demand of 0.25 kWh/m3STP including biogas compression [232, p. 61].

216


A.2.1.12 Gas turbine system

Table A.24: Modelling assumptions for gas turbine systems, and resulting efficiencies fornatural gas-fuelled operation.

large medium

capacity [MWel] >300 30–70

combustor outlet temperature [°C] 1350 1230

compressor pressure ratio [–] 15.38 15.38

compressor polytropic efficiency [–] 92.0% 90%

expander isentropic efficiency (per stage) [–] 91.6% 87.8%

compressor & expander mechanical

efficiency

[–] 99% 99%

generator efficiency [–] 98.5% 98.5%

combustor heat loss [% fuel HHV] 1.5% 1.5%

combustor pressure loss [% ] 3% 3%

cooling air demand [kgair/kggas] 0.106 0.1233

distribution of cooling air between stages [–] 62% / 29% / 9% / 0% 60% / 30% / 10% / 0%

calculated efficiency (HHV) [–] 35.2% 31.8%

calculated efficiency (LHV) [–] 39.0% 35.3%

A.2.1.13 Air separation unit

Table A.25: Modelling assumptions for the air separation units.

integrated standaloneinlet pressure [bar] 15.4 4.6outlet pressure [bar] 4 1.0oxygen purity [mol% O2] 95% 95%oxygen recovery rate [–] 92.5% 92.5%cold box electricity consumption [kJ/kgO2] 100 100cold box LP steam consumption [kJ/kgO2] 163 163

A.2.1.14 Reciprocating engine

The electric output is calculated with the given electrical efficiency from a technical datasheet [377]. The engine cooling water duty is then calculated by energy balance based onthe enthalpy of reaction of the biogas combustion and the given exhaust gas temperature.The calculated HHV-based efficiency is 38.3%.

Table A.26: Modelling assumptions for the reciprocating engine.

electrical efficiency, LHV [–] 42.5% [377]exhaust gas temperature [°C] 454 [377]air ratio [–] 1.5heat loss [% of fuel, HHV] 3%

217


A.2.1.15 Miscellaneous auxiliary units

In the upgrading plants, waste heat is discharged to the environment via air coolers. Thecooling air is not explicitly modelled in the Aspen Plus simulations. The cooling air volumeflow is calculated on the assumption that ambient air is heated from 15 to 30°C, with anaverage heat capacity of 1.005 kJ/kg/K and a density of 1.216 kg/m3. The electricityconsumption of the cooling air fans is assumed to be 0.50 kJ/m3 cooling air.

Cooling water for the condenser and intercooled compressors in the IGCC plants is notexplicitly modelled in the Aspen Plus simulations. The required cooling water demand iscalculated assuming a temperature increase of 10°C in the cooling water with an averageheat capacity of 4.19 kJ/kg/K. The electricity consumption of the cooling water pumps isassumed to be 0.061 kJ/kg.

Other miscellaneous electricity consumers in the IGCC plants (such as water demineral-ization, waste water treatment, slag treatment, control systems, lighting) are assumed todemand 0.002 MJ/MJ of the gross electricity production.

A.3 Assumptions for the cost estimates

A.3.1 Adjustment to year 2010 € and plant location

Costs can be adjusted to different geographical locations with the help of location factors,which take into account regional conditions such as the cost and productivity of labour,equipment, commodities and taxes, as well as exchange rates [378, p. 259]. Cost datagiven in US dollars (USD) from year y is converted to year 2010 Euro as suggested by[304, p. 100], with a location factor of 1.11 for Germany compared to the USA in 2003[379, p. 327] and an average annual exchange rate in 2003 of 0.885 €/USD4:

C€,2010 = 1.11 · 0.885€2003

USD2003· CEPCI2003

CEPCIy· D4102010

D4102003· CUSD,y

The location factor and exchange rate result in a conversion factor from USD to Euro of0.9827 in 2003. The Chemical Engineering Plant Cost Index (CEPCI) [380] is applied toconvert cost data given in USD to year 2003 USD, and an index for chemical engineeringequipment published by the Federal Statistical Office of Germany (D410) [381] is used toadjust cost data in € to the year 2010.

Location factors for Malaysia and South-East Asia are, in general, similar to those forGermany [382, p. 187] [379, p. 327], therefore investment cost estimates for a Germanlocation are used for Malaysia without modification.

A.3.2 Equipment cost CBM

The cost data used for calculating the module cost (CBM) of the plant equipment is listedin Table A.27 to Table A.45. Modifications of these functions are made due to special

4PACIFIC Exchange Rate Service, yearly average.

218


design requirements in the individual plant designs, for example because special materialsare required. These modifications are discussed in the appendix section Cost data for therespective plant models.

If n characteristic capacities Xi, ref are given, equation Equation 3.11 is modified to:

C = Cref ·(

a1

(X1

X1,ref

)α

+ ... + an

(Xn

Xn,ref

)α)

where the weighting factor ai is 1/n unless otherwise stated. This is based on the assump-tion that, for example for an AGR unit which removes CO2 and H2S, some parts of theunit are sized based on the required capacity for H2S removal, while others are sized basedon the required CO2 removal. For cost functions, the range of validity for the capacity Xis given.

If the year of the cost data is not stated in the source, it is assumed to be 1–2 years prior topublication. A currency in brackets indicates the currency of the original source, when thecost function has been converted to € 2010 (in order to combine it with data from anothercost source to adjust the scope of supply). Cref is given in Table A.27 to Table A.45 inthe currency as stated in the source.

(PEC) indicates that Cref or α is based on the purchased equipment cost. Cref then has tobe converted to CBM by multiplication with CBM/PEC. If no component specific datawas available, CBM/PEC is assumed to be 2.46 for components processing solids, 3.32for components processing fluids, and 2.88 for components processing a mixture of both,based on data from [306, p. 252]. Conversion of α to CBM basis is performed by weightingαP EC and αother = 0.64.

Values in brackets for α are used when the cost function has to be used outside its recom-mended range. In this case (which is avoided where possible), the original cost functionis applied up to the limit of its validity Xmin or Xmax, and scaling below Xmin or aboveXmax is undertaken with equation Equation 3.11 using C(Xmin) or C(Xmax) as Cref.

If no values were found in literature, α is estimated by the author.

The material factor fM for carbon steel (CS) is 1.0 unless otherwise stated.

fp = pressure factor

fSH = temperature factor per °C of superheating

219


Tabl

eA

.27:

Cos

tda

tafo

rhe

atex

chan

gers

.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

shel

l&tu

beH

X-1

10-1

000

m2

USD

2003

f(A

,p)

(0.5

9)SS

/CS=

1.81

SS/S

S=2.

73

[307

]

shel

l&tu

beH

X-2

1000

m2

GB

P20

0012

01.

80.

662

see

HX

-1[3

04,F

ig.

7.18

]

HR

SGH

XH

X-3

[kW

/K]

USD

1985

PE

C(k

A)

1.8

SH=

2.0

CG

=2.

0

[383

]

syng

asH

XH

X-4

6700

m2

€20

1037

540.

6[3

84]

spir

alpl

ate

HX

S-1

10-1

00m

2U

SD20

03f(

A,p

)se

eH

X-1

see

HX

-1[3

07]

air-

cool

erA

CO

-110

-100

00m

2U

SD20

03f(

A,p

)se

eH

X-1

SS=

2.93

[307

]

rota

ryai

rpr

ehea

ter

AP

H-1

[m2]

USD

1990

f(A

)1.

50.

6(P

EC

)[3

85]

stea

mtu

rbin

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nden

ser

CN

D-1

372

MW

thU

SD20

0766

200.

9[2

46]

Com

men

ts:

HX

-1:

Fix

edtu

be,s

heet

,or

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be.α:

shel

l&tu

beH

X[3

07,p

.14

9]

HX

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carb

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eels

hell

&tu

beH

X.α

deri

ved

for

1000

-700

00m

2[3

04,F

ig.

7.18

]

HX

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cost

func

tion

:P

EC

=2.

975(

kA

)0.8

[383

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ovid

esco

stfu

ncti

ons

for

the

indi

vidu

alco

mpo

nent

sof

aH

RSG

(ind

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duct

s,ca

se&

stac

k).

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ause

they

are

base

don

old

data

(bef

ore

1986

),th

eir

valid

ity

was

test

edby

appl

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mul

atio

nda

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ra

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SGfr

om[3

09,c

ase

3A].

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sum

ofP

EC

calc

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the

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func

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sfr

om[3

83]i

s17

%hi

gher

than

the

PE

Cfo

rth

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eral

lHR

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ven

in[3

09],

whi

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cons

ider

eto

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cept

able

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BM

/PE

C=

2.0

incl

udin

gst

art-

upco

st.

Bas

edon

data

from

[309

],C

BM

/PE

C=

1.5

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=su

perh

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cond

ensi

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haus

tga

s

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for

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omen

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flow

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omas

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urry

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85],

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don

[386

]

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epe

lleti

zing

equi

pmen

t

220


Tabl

eA

.28:

Cos

tda

tafo

rtu

rbin

es,p

umps

,com

pres

sors

,fan

san

dm

otor

s.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

turb

ines

stea

mtu

rbin

e,C

HP

STB

-10.

5-10

MW

el€

2000

f(W

)[3

87,p

.II

.79]

stea

mtu

rbin

e,C

HP

STB

-20.

5-10

MW

el€

2000

f(W

)[3

87,p

.II

.79]

stea

mtu

rbin

e,IG

CC

STC

-120

9.4

MW

elU

SD20

0733

621

0.9

[311

,p.

343]

[246

]

syng

asex

pand

erSG

E-1

10M

Wel

USD

2003

2892

(PE

C)

1.2

0.67

[388

]

pu

mp

s

pum

pP

-11-

300

kWU

SD20

03f(

W)

(0.7

1)(P

EC

)

(0.6

64)(

CB

M)

CS=

1.5

SS=

2.2

[307

]

com

pre

ssor

s&

fan

s

com

pres

sor,

smal

lC

MP

-10.

5-3

MW

USD

2003

f(W

)(0

.84)

CS=

2.7

SS=

5.8

[307

]

com

pres

sor,

smal

lC

MP

-218

-950

kWU

SD20

03f(

W)

(0.8

4)C

S=2.

4

SS=

5.0

[307

]

air

orN

2co

mpr

esso

rC

MP

-310

MW

elU

SD20

0349

68(P

EC

)1.

20.

67[3

55]

syng

asco

mpr

esso

rC

MP

-410

MW

elU

SD20

0357

96(P

EC

)1.

20.

67[3

55]

O2

com

pres

sor

CM

P-5

10M

Wel

USD

2003

6648

(PE

C)

1.2

0.67

[355

]

CO

2co

mpr

esso

rC

MP

-67.

55M

Wel

USD

2007

7766

0.67

[246

]

fan

AF

-11-

100

m3/s

USD

2003

f(V

,p)

(0.6

)[3

07]

mot

ors

elec

tric

alm

otor

EM

-10.

07-2

.6M

WU

SD20

03f(

W)

(0.8

4)[3

07]

Com

men

ts:

STB

-1,

STB

-2:

indu

stri

alba

ckpr

essu

rest

eam

turb

ine,

incl

udin

gge

nera

tor

STC

-1:

larg

eco

nden

sing

stea

mtu

rbin

e,in

clud

ing

stea

mpi

ping

and

auxi

liari

es

SGE

-1:

valu

egi

ven

for

PE

Cis

inst

alle

dco

st(I

C),

CB

M/P

EC

=1.

2(I

C+

20%

for

engi

neer

ing)

.P

EC

andα

from

onlin

esu

pple

men

tary

mat

eria

lfor

[388

]

P-1

:ce

ntri

fuga

lpum

p,α P

EC

from

[385

].Sl

urry

pum

pfo

rH

TC

plan

t:se

ebi

omas

str

ansp

ort

and

pres

suri

zati

on

221


CM

P-1

:ce

ntri

fuga

lcom

pres

sor,

wit

hout

mot

or.α

for

reci

proc

atin

gco

mpr

esso

r[3

07,p

.14

9]

CM

P-2

:ro

tary

com

pres

sor,

wit

hout

mot

or.α

for

reci

proc

atin

gco

mpr

esso

r[3

07,p

.14

9]

CM

P-3

,C

MP

-4,

CM

P-5

,:

valu

egi

ven

for

PE

Cis

inst

alle

dco

st,C

BM

/PE

C=

+20

%fo

ren

gine

erin

g.α

from

[355

]

CM

P-6

:30

.21

MW

,31.

1M

USD

,4un

its.

Cos

tis

give

nhe

refo

r1

unit

.α

from

[355

]

AF

-1:

axia

ltub

efa

n,in

clud

ing

mot

or

EM

-1:

open

/dri

p-pr

oof

elec

tric

driv

e

Tabl

eA

.29:

Cos

tda

tafo

rbo

ilers

and

furn

aces

.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

woo

d-fir

edfu

rnac

eW

F-1

3704

kWH

HV

USD

2004

214

0.66

[112

]

woo

d-fir

edbo

iler,

smal

lW

B-1

1kg

stea

m/s

USD

2004

144

0.79

6[3

05,F

ig.

5.4]

woo

d-fir

edbo

iler,

larg

eW

B-2

87M

WH

HV

USD

2005

6389

0.78

1[3

14]

gas

boile

rG

B-1

1kg

stea

m/s

USD

2004

900.

815

[305

,Fig

.5.

4]

offga

sin

cine

rato

rO

GI-

11

MW

USD

2004

500

0.59

4[3

05,F

ig.

5.10

]

Com

men

ts:

WF

-1:

solid

fuel

furn

ace

for

woo

dsh

avin

gs,

w=

10%

.T

hefir

ing

capa

city

ises

tim

ated

base

don

the

dem

and

for

dryi

ng6

t/h

woo

dch

ips

from

40%

to10

%w

ater

cont

ent,

assu

min

ga

spec

ific

heat

dem

and

of40

00kJ

/kg H

2O,e

van

da

ther

mal

effici

ency

of90

%of

the

furn

ace.

WB

-1:

pack

aged

,coa

l-fire

dbo

iler.α

and

f pfo

rp=

20ba

rde

rive

dfr

om[3

05]

WB

-2:

woo

d-fir

edst

oker

boile

r.α

and

f pba

sed

on[3

14],

f SH

base

don

[305

,Fig

.5.

4]

GB

-1:

oil-

and

gas-

fired

boile

r.α

and

f pfo

rp=

20ba

rde

rive

dfr

om[3

05]

OG

I-1:α

deri

ved

for

data

from

1–5

MW

,bas

edon

[305

,Fig

.5.

10]

222


Tabl

eA

.30:

Cos

tda

tafo

rdr

iers

.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αso

urce

s

rota

rydr

umdr

ier

RD

D-1

6400

kgH

2O,e

v/h

€20

1095

0(P

EC

)

1250

(CB

M)

1.32

0.50

(PE

C)

[327

],

0.58

3(C

BM

)

see

com

men

ts

rota

rydr

umdr

ier

RD

D-2

6000

kgH

2O,e

v/h

€20

1010

00se

eR

DD

-1se

eco

mm

ents

rota

rydr

umdr

ier

RD

D-3

8000

kgH

2O,e

v/h

€20

1020

00se

eR

DD

-1se

eco

mm

ents

belt

drie

rB

D-1

833

kgH

2O,e

v/h

€20

0670

0(P

EC

)se

eR

DD

-1se

eR

DD

-1[3

89,p

.38

]

belt

drie

rB

D-2

3000

kgH

2O,e

v/h

€20

0250

0(P

EC

)se

eR

DD

-1se

eR

DD

-1[1

06]

SSD

drie

rSS

D-1

3000

kgH

2O,e

v/h

€20

0214

60(P

EC

)se

eR

DD

-1se

eR

DD

-1[1

06]

Com

men

ts:

RD

D-1

:na

tura

lgas

fired

drum

drie

rfo

rdr

ying

woo

dch

ips

from

50%

to10

%w

ater

cont

ent,

hot

gas

at45

0°C

,exh

aust

gas

110°

C.C

BM

=P

EC

+st

ruct

ural

stee

l+in

sula

tion

,ve

ndor

s’in

form

atio

nci

ted

in[1

43].

The

scop

eof

supp

lyin

clud

esa

natu

ral

gas

burn

er,

whi

chis

not

requ

ired

for

the

cons

ider

edpl

ant

desi

gn,

beca

use

cost

sfo

rth

ew

ood

furn

ace

are

acco

unte

dfo

rse

para

tely

.P

roba

bly

addi

tion

alga

scl

eani

ngto

redu

cedu

stan

dV

OC

emis

sion

sis

requ

ired

.It

isas

sum

edth

atth

eco

stre

duct

ion

for

omit

ting

the

gas

burn

eran

dth

eco

stin

crea

sedu

eto

exha

ust

gas

clea

ning

com

pens

ate

for

each

othe

r.

RD

D-2

:na

tura

lgas

fired

drum

drie

rfo

rw

ood

chip

s,in

clud

ing

burn

er.

Ven

dors

’inf

orm

atio

nci

ted

in[1

43].

Com

men

ton

scop

eof

supp

ly:

see

RD

D-1

RD

D-3

:dr

umdr

ier

for

woo

dch

ips,

incl

udin

gbu

rner

,fee

ding

bin,

cycl

ones

,sta

ck,i

nstr

umen

tati

on,s

truc

tura

lste

el,v

endo

rs’i

nfor

mat

ion

cite

din

[143

].C

omm

ent

onsc

ope

ofsu

pply

:se

eR

DD

-1.

BD

-1:

belt

drie

rfo

rbi

omas

s.C

apac

ity

calc

ulat

edfr

om1

MW

th,

spec

ific

heat

dem

and

4320

kJ/k

gH2O

,ev.

Max

.ca

paci

type

run

it:

44t H

2O,e

v/h

,ba

sed

onre

aliz

edca

paci

ties

of1-

50t/

hpr

oduc

tfo

rlo

wte

mpe

ratu

resa

wdu

stbe

ltdr

iers

[390

]

BD

-2:

belt

drie

rfo

rdr

ying

woo

dch

ips

from

55%

to10

%w

ater

cont

ent,

oper

atin

gpr

essu

re0.

5-2.

0ba

r

SSD

-1:

SSD

drie

rfo

rdr

ying

woo

dch

ips

from

55%

to10

%w

ater

cont

ent,

oper

atin

gpr

essu

re0.

5-2.

0ba

r

223


Tabl

eA

.31:

Cos

tda

tafo

rm

echa

nica

ldew

ater

ing

equi

pmen

t.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

filte

rpr

ess

FP

-110

m3/h

€20

0514

00.

706

SS=

1.5

[374

]

filte

rpr

ess

FP

-28.

8m

3/h

€20

0133

20.

706

SS=

1.5

[33]

deca

nter

DE

C-1

0.06

kgD

M/s

USD

2004

600.

365

CS=

2.0

SS=

3.4

[305

,Fig

.5.

54]

scre

wpr

ess

SWP

-10.

402

kgD

M/s

USD

2004

700.

546

CS=

2.4

SS=

3.6

[305

,Fig

.5.

58]

Com

men

ts:

FP

-1:

cham

ber

filte

rpr

ess

for

sew

age

slud

ge,

dew

ater

ing

from

5%to

35%

dry

mat

ter.

f p=

1.2

for

high

pres

sure

filte

rpr

ess,

base

don

cost

data

for

2pl

ants

[374

].f M

:ra

tio

betw

een

SSan

dC

Sfo

rfil

ter

pres

sba

sed

on[3

05,

Fig

.5.

57b]

.Si

nce

the

bioc

oal

isea

sier

tode

wat

erth

anse

wag

esl

udge

,it

isas

sum

edth

atth

eco

stfo

ra

filte

rpr

ess

dew

ater

ing

bioc

oalt

o60

%dr

ym

atte

ris

the

sam

eas

for

dew

ater

ing

sew

age

slud

geto

abou

t30

%dr

ym

atte

r.

FP

-2:

cham

ber

filte

rpr

ess

for

sew

age

slud

ge,d

ewat

erin

gfr

om3%

to25

%dr

ym

atte

r.f p

and

f M:

see

FP

-1

DE

C-1

:α

deri

ved

for

0.06

-0.5

kg/s

from

[305

,Fig

.5.

54]

SWP

-1:α

deri

ved

for

0.1-

1.0

kg/s

from

[305

,Fig

.5.

58]

Tabl

eA

.32:

Cos

tda

tafo

rsc

reen

ing

and

sizi

ngeq

uipm

ent.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αso

urce

s

shre

dder

&dr

umsc

reen

BSS

-110

.8t/

h€

2001

136

0.67

2[3

3,T

ab.

A.6

.2]

[33]

mag

neti

cse

para

tor

BSS

-210

.8t/

h€

2001

150.

5[3

3,T

ab.

A.6

.2]

air

clas

sifie

rB

SS-3

10.8

t/h

€20

0118

0.5

[33,

Tab

.A

.6.2

]

ham

mer

mill

HM

-155

00kg

/h€

2010

133

1.5

0.66

5se

eco

mm

ents

Com

men

ts:

BSS

-1:

for

gard

enin

gw

aste

,com

post

,bar

km

ulch

,was

tew

ood.α

deri

ved

for

data

from

3.6

and

10.8

t/h

plan

ts.

BSS

-2,

BSS

-3:

for

gard

enin

gw

aste

,com

post

,bar

km

ulch

,was

tew

ood

224


HM

-1:

ham

mer

mill

for

woo

dch

ips

wit

h6

mm

scre

en,v

endo

rs’i

nfor

mat

ion

cite

din

[143

]

coal

mill

atIG

CC

:inc

lude

din

feed

ing

syst

emfo

rE

Fga

sifie

r,se

ebi

omas

str

ansp

ort

&pr

essu

riza

tion

Tabl

eA

.33:

Cos

tda

tafo

rfe

edin

gan

dpr

essu

riza

tion

equi

pmen

t.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αso

urce

s

slur

rypu

mp

(HT

C)

SLP

-110

m3/h

€20

1017

70.

664

(see

pum

pP

-1)

[336

,p.5

5-57

]

plug

-for

min

gfe

eder

PSF

-141

.7t/

h€

2010

1000

0.63

1[3

36,p

.43

]

scre

wco

nvey

orSC

C-1

5.5

kg/s

USD

2002

20(P

EC

)2.

460.

418

(PE

C)

[308

,Fig

.

12-6

0]

feed

ing

syst

emfo

rE

Fga

sifie

rE

FF

-129

.3kg

/sU

SD20

0774

430

0.75

[246

]

feed

ing

syst

emfo

rF

Bga

sifie

rF

BF

-117

.94

kg/s

USD

1999

788

0.77

[355

][3

55]

feed

ing

syst

emfo

rF

Bga

sifie

rF

BF

-23.

83kg

/sU

SD20

0038

01.

0[2

97]

[297

]

Com

men

ts:

SLP

-1:

pist

onpu

mp

for

high

-den

sity

solid

s,Δ

p=30

bar.

Cos

tda

taba

sed

ona

vend

ors

quot

efo

ran

HT

Cap

plic

atio

n,ci

ted

in[3

36,

p.55

-57]

.Fo

rde

sign

sw

ith

mix

ing

preh

eate

rsbe

twee

npu

mpi

ngst

ages

,C

BM

isas

sum

edto

be+

20%

for

2st

ages

and

+30

%fo

r4

stag

es.

70%

ofth

eC

BM

are

assu

med

tobe

depe

nden

ton

pres

sure

head

:C

BM

=C

BM

30ba

r

( 0.7

Δp

30ba

r+

0.3)

PSF

-1:

Cos

tda

taba

sed

onve

ndor

sin

form

atio

nfo

ra

plug

scre

wfe

eder

,ca

paci

ty10

00t/

d,pr

essu

re15

bar(

g).

Acc

ordi

ngto

the

vend

or,

high

erpr

essu

res

can

beac

hiev

edby

exte

ndin

gth

epl

ugsc

rew

.T

hefo

llow

ing

assu

mpt

ion

ism

ade

for

adju

stin

gth

eco

stto

the

high

erpr

essu

re:

CB

M=

CB

M15

ba

r

( 0.7

Δp

15ba

r+

0.3)

SCC

-1:

scre

wco

nvey

orfo

rgr

ain,

l=30

m,d

=0.

23.

leng

thfa

ctor

:f

l=

( l 30m

) 0.678.

The

leng

thfo

rtr

ansp

orti

ngfe

edst

ock

toth

eto

pof

the

HT

Cre

acto

ris

calc

ulat

edw

ith

anan

gle

ofde

clin

eof

15°.α

and

PE

Cde

rive

dfr

om[3

08,F

ig.

12-6

0]fo

r5.

5-27

.5kg

/san

dl=

4-30

m

EF

F-1

:In

clud

eslo

ckho

pper

,coa

lmill

,coa

ldri

erfo

rdr

ying

from

10%

to5%

wat

erco

nten

t.2

trai

nsw

ith

ato

talc

apac

ity

of21

1t/

hco

st14

9M

USD

.Cos

tar

egi

ven

here

for

1tr

ain

FB

F-1

:co

mpr

isin

gfe

ed-b

in,3

0ba

rai

r-lo

ck,f

eed-

scre

w,v

alue

sfo

rpr

oces

sing

woo

d,ba

sed

onda

tafr

omW

eyer

heus

erco

mpa

ny,2

000

FB

F-2

:fe

edin

gsy

stem

for

69.5

4M

WL

HV

biom

ass

gasi

fier

225


Tabl

eA

.34:

Cos

tda

tafo

rso

lids

hand

ling

equi

pmen

tat

IGC

Cpl

ants

.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

coal

,p

elle

tsfr

omH

TC

,

TO

R

coal

hand

ling

CH

-158

.6kg

/sU

SD20

0729

700

0.7

[246

]

woo

dch

ips

woo

dco

nvey

orW

C-1

17.9

4kg

/sU

SD19

9985

10.

77[3

55]

[355

]

woo

dco

nvey

orW

C-2

5.47

kg/s

USD

2000

430

0.8

[297

][2

97]

woo

dst

orag

eW

S-1

17.9

4kg

/sU

SD19

9956

10.

77[3

55]

[355

]

woo

dst

orag

eW

S-2

4.26

kg/s

USD

2000

1050

0.65

[297

][2

97]

iron

rem

oval

IR-1

4.26

kg/s

USD

2000

330

0.7

[297

][2

97]

ash

,sl

ag

slag

hand

ling

SH-1

0.81

5kg

/sU

SD20

0716

000

0.7

[384

][2

46]

ash

hand

ling

AH

-10.

751

kg/s

USD

2007

7200

0.7

[384

][2

46,c

ase

10]

Com

men

ts:

WC

-1:

valu

esfo

rpr

oces

sing

woo

d,ba

sed

onda

tafr

omW

eyer

heus

erco

mpa

ny,2

000

WC

-2:

feed

ing

syst

emfo

r69

.54

MW

LH

Vbi

omas

sga

sifie

r,w

ood

chip

sw

ith

30%

wat

erco

nten

t

WS-

1:dr

ied

woo

dch

ips

stor

age,

base

don

data

from

Wey

erhe

user

com

pany

,200

0

WS-

2,IR

-1:

stor

age

for

69.5

4M

WL

HV

biom

ass

gasi

fier,

drie

dw

ood

chip

sw

ith

10%

wat

erco

nten

t

SH-1

:lo

ckho

pper

for

depr

essu

riza

tion

,cru

sher

,sto

rage

,21.

1t/

h,32

MU

SD,2

unit

s.C

ost

isgi

ven

here

per

unit

AH

-1:

ash

hand

ling

syst

emfo

rpu

lver

ized

coal

-fire

dpo

wer

plan

t,in

clud

ing

hopp

ers,

stor

age,

5.41

t/h,

14.3

MU

SD,2

unit

s.C

ost

isgi

ven

here

per

unit

Whe

nw

ood

pelle

tsin

stea

dof

woo

dch

ips

are

proc

esse

d,th

eco

stfo

rth

ew

ood

conv

eyor

and

stor

age

ism

ulti

plie

dby

0.5,

iron

rem

oval

isno

tre

quir

ed.

226


Tabl

eA

.35:

Cos

tda

tafo

rpe

lleti

zing

and

pelle

tha

ndlin

geq

uipm

ent.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

pelle

tpr

ess

PP

-130

00kg

/h€

2010

401

0.71

see

com

men

ts

pelle

tpr

ess

PP

-250

00kg

/hC

AD

2010

771

0.71

see

com

men

ts

pelle

tco

oler

PC

-130

00kg

/h€

2010

158

0.59

(PE

C)

0.62

(CB

M)

see

com

men

ts

pelle

tsi

evin

gan

dco

ntro

l

scre

enin

g

PSS

-130

00kg

/h€

2010

310.

6se

eco

mm

ents

pelle

tst

orag

eP

S-1

5000

kg/h

CA

D20

1056

40.

66se

eco

mm

ents

Com

men

ts:

All

data

inT

able

A.3

5is

base

don

vend

ors’

info

rmat

ion

cite

din

[143

]

PP

-1,

PP

-2:

incl

udin

gco

ndit

ioni

ng

PC

-1:

air-

cool

ed

Tabl

eA

.36:

Cos

tda

tafo

rta

nks.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

resi

due

stor

age

tank

RST

-111

00-1

1400

m3

€20

07f(

V)

1.2

[344

]

tank

TN

K-1

0.1-

628

m3

USD

2003

f(V

,d,p

)SS

-cla

d=1.

7[3

07]

tank

TN

K-2

10-1

00m

3U

SD20

04f(

V)

3.32

(0.6

)SS

-cla

d=2.

5[3

06,p

.25

9]

Com

men

ts:

RST

-1:

cost

func

tion

:c=

198.

71V

^(-0

.154

2)w

here

cis

spec

ific

PE

Cin

[€20

07/m

3],

and

Vis

the

volu

me

in[m

3],

deri

ved

from

data

for

9pl

ants

base

don

[344

]

TN

K-1

:ho

rizo

ntal

vess

el

TN

K-2

:ho

rizo

ntal

tank

,fp

and

f Mfr

om[3

05,F

ig.

5.45

]

227


Tabl

eA

.37:

Cos

tda

tafo

rre

acto

rs.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

torr

efac

tion

reac

tor

TO

R-1

9574

kgfe

ed/h

€20

0355

02.

460.

72[3

27].

[114

]

HT

Cre

acto

rH

TC

-10.

3-52

0m

3U

SD20

03f(

V,d

,p)

SS-c

lad=

1.7

[307

]

HT

Cre

acto

rH

TC

-21-

50m

3U

SD20

04f(

V)

3.32

(0.6

)SS

-cla

d=2.

5[3

06,p

.25

9]

HT

Cre

acto

rH

TC

-318

.9m

3U

SD19

9830

(PE

C)

3.17

0.6

[306

]SS

=2.

8[3

91,p

.6]

AD

dige

ster

AD

F-1

620-

2000

m3

€20

07f(

V)

1.9

[232

,344

]

[344

]

stir

rer

STI-

145

kWU

SD20

0442

0.80

4SS

=2.

5[3

05,F

ig.

5.42

]

Com

men

ts:

TO

R-1

:m

ovin

gbe

dre

acto

r,pr

oduc

tion

capa

city

of60

ktto

rrefi

edw

ood

HT

C-1

:ve

rtic

alve

ssel

HT

C-2

:ve

rtic

alta

nk,

f pan

df M

from

[305

,Fig

.5.

45]

HT

C-3

:ve

rtic

alve

ssel

,p=

10ba

r.f M

for

stai

nles

sst

eel

type

304

[391

,T

ab.

7],

f pfr

om[3

05,

Fig

.5.

45],

CB

M/P

EC

:bu

lkin

stal

lati

ons

for

liqui

dan

dsl

urry

syst

ems

wit

hp>

10ba

r[3

91,T

ab.

4]

AD

F-1

:co

stfu

ncti

on:

c=10

38.9

V^(

-0.3

424)

whe

rec

issp

ecifi

cP

EC

in[€

2007

/m3],

and

Vis

the

volu

me

in[m

3],

deri

ved

from

data

for

9pl

ants

base

don

[344

]

STI-

1:ag

itat

or,o

pen

tank

.α

deri

ved

for

10-1

00kW

from

[305

,Fig

.5.

42]

228


Tabl

eA

.38:

Cos

tda

tafo

rga

sifie

rs.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

entr

aine

dflo

wga

sifie

rE

FG

-139

.6€

2010

(USD

2007

)

1315

720.

8[3

84]

[246

]

entr

aine

dflo

wga

sifie

rE

FG

-239

.6€

2010

(USD

2001

)

1187

130.

8[3

84]

[256

,p.

18].

fluid

ized

bed

gasi

fier

FB

G-1

11.5

8kg

DM

/sU

SD20

0364

100.

7[3

55]

[355

]

fluid

ized

bed

gasi

fier

FB

G-2

11.3

2kg

DM

/sU

SD20

0387

000.

7[3

92]

fluid

ized

bed

gasi

fier

FB

G-3

7.12

kgD

M/s

€20

0916

100

0.7

[256

]

fluid

ized

bed

gasi

fier

FB

G-4

23.1

3kg

DM

/s€

2010

(USD

2008

)

1937

00.

7[3

93].

fluid

ized

bed

gasi

fier

FB

G-5

20.4

6kg

DM

/sU

SD20

0030

000

0.7

[297

][2

97]

Com

men

ts:

EF

G-1

:Sh

ell

coal

gasi

fier,

197.

4t/

hfe

ed,

2un

its,

incl

.sy

ngas

cool

er,

cycl

one,

cand

lefil

ter,

foun

dati

ons,

flare

stac

k,24

1.5

MU

SD20

07.

Thi

sre

sult

sin

134

M€

2010

for

a39

.6kg

/sun

itfr

oma

sim

ulat

ion

for

bitu

min

ous

coal

gasi

ficat

ion,

ofw

hich

the

cost

for

2sy

ngas

cool

ers

(154

MW

,50

10m

2on

tota

l,co

stca

lcul

ated

wit

hH

X-1

,H

X-2

,H

X-4

)ar

esu

btra

cted

.

EF

G-2

:Sh

ellc

oalg

asifi

er,2

200

t DM

/dfe

ed,i

ncl.

syng

asco

olin

gto

240°

C(p

.18

),8

0.5

MU

SD20

01.

Thi

sre

sult

sin

122

M€

2010

for

a39

.6kg

/sun

itfr

oma

sim

ulat

ion

for

bitu

min

ous

coal

gasi

ficat

ion,

ofw

hich

the

cost

for

2sy

ngas

cool

ers

(154

MW

,501

0m2

onto

tal,

cost

calc

ulat

edw

ith

HX

-1,H

X-2

,HX

-4)

are

subt

ract

ed.

FB

G-1

:pr

essu

rize

d,ox

ygen

-blo

wn

biom

ass

gasi

fier,

30ba

r,10

14°C

FB

G-2

:pr

essu

rize

d,st

eam

and

oxyg

en-b

low

nbi

omas

sga

sifie

r,20

bar,

850°

C

FB

G-3

:bu

bblin

gflu

idiz

edbe

dbi

omas

sga

sifie

r

FB

G-4

:pr

essu

rize

d,ox

ygen

-blo

wn

fluid

ized

-bed

gasi

fier,

389

MW

LH

Vco

rnst

over

(w=

10%

),30

bar,

870°

C,2

8.2

MU

SD20

08in

clud

ing

lock

hopp

ers.

Cos

tfo

rlo

ckho

pper

ssu

btra

cted

(1.9

9M

€20

10,b

ased

onF

BF

-1,F

BF

-2)

FB

G-5

:pr

essu

rize

d,ci

rcul

atin

gflu

idiz

edbe

dIG

Tbi

omas

sga

sifie

r,ox

ygen

-blo

wn,

400

MW

HH

V,3

0ba

r,10

14°C

229


Tabl

eA

.39:

Cos

tda

tafo

rpa

rtic

ulat

ere

mov

alan

dm

isce

llane

ous

gas

clea

ning

equi

pmen

t.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

syng

assc

rubb

erSG

S-1

20.9

m3/s

USD

2003

660

(PE

C)

1.2

0.7

[355

][3

55]

cycl

one

CY

C-1

68.7

m3/s

USD

2003

910

(PE

C)

1.2

0.7

[355

][3

55]

cycl

one

CY

C-2

34.2

m3/s

€20

0660

000.

7[3

55]

[394

]

cand

lefil

ter

CF

-114

.4m

3/s

USD

2003

1860

0(P

EC

)1.

20.

65[3

55]

[355

]

cand

lefil

ter

CF

-22.

71m

3/s

€20

0997

000.

65[3

55]

[392

]

bagh

ouse

filte

rC

F-3

0.89

5m

3/s

USD

2000

1620

0.7

[297

][2

97]

syng

assa

tura

tor

SAT

-120

.9m

3/s

USD

2003

220

(PE

C)

1.2

0.7

[355

][3

55]

mis

c.sy

ngas

trea

tmen

tM

ST-1

4.00

5km

ol/s

USD

2007

4328

0.7

[246

]

Com

men

ts:

SGS-

1:va

lue

give

nfo

rP

EC

isin

stal

led

cost

(IC

),C

BM

/PE

C=

1.2

(IC

+20

%fo

ren

gine

erin

g)

CY

C-1

:cy

clon

efo

rbi

omas

sga

sifie

r,va

lue

give

nfo

rP

EC

isin

stal

led

cost

(IC

),C

BM

/PE

C=

1.2

(IC

+20

%fo

ren

gine

erin

g)

CY

C-2

:cy

clon

efo

rbi

omas

sga

sifie

r

CF

-1:

cera

mic

cand

lefil

ter

for

biom

ass

gasi

fier,

valu

egi

ven

for

PE

Cis

inst

alle

dco

st(I

C),

CB

M/P

EC

=1.

2(I

C+

20%

for

engi

neer

ing)

CF

-2:

cera

mic

cand

lefil

ter

for

biom

ass

gasi

fier

wit

h46

7M

WH

HV

,Vg

as/Q

fe

ed

,HH

Vfr

omsi

mul

atio

nF

B-w

ood-

1

CF

-3:

bagh

ouse

filte

rfo

rbi

omas

sga

sifie

rw

ith

427

MW

HH

V,V

ga

s/Q

fe

ed

,HH

Vfr

omsi

mul

atio

nF

B-w

ood-

1

SAT

-1:

valu

egi

ven

for

PE

Cis

inst

alle

dco

st(I

C),

CB

M/P

EC

=1.

2(I

C+

20%

for

engi

neer

ing)

MST

-1:

mer

cury

rem

oval

,blo

wba

cksy

stem

,fue

lgas

pipi

ng,f

ound

atio

ns,2

unit

s.C

ost

are

give

nhe

repe

run

it.

Xre

fis

syng

asflo

wat

shift

reac

tor

outl

et

230


Tabl

eA

.40:

Cos

tda

tafo

rde

sulp

huri

zati

onan

dac

idga

sre

mov

aleq

uipm

ent.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

αso

urce

s

pres

suri

zed

wat

ersc

rubb

ing

PW

S-1

500

m3/h

raw

gas

€20

0713

240.

360

[232

]

acid

gas

rem

oval

for

CO

2an

dH

2S

AG

R-1

8.12

kmol

gas/

s,

3.07

kmol

CO

2/s

,

0.04

6km

olH

2S/s

USD

2007

1422

740.

65[2

46]

acid

gas

rem

oval

for

CO

2an

dH

2S

AG

R-2

10.2

8km

olga

s/s,

3.81

kmol

CO

2/s

,

0.01

9km

olH

2S/s

USD

2001

8552

60.

65[3

00,c

ase

B1]

acid

gas

rem

oval

for

CO

2on

lyA

GR

-32.

75km

olC

O2/s

€20

0663

000

0.65

[394

]

sulp

hur

reco

very

SRU

-11.

47kg

sulp

hur/

sU

SD20

0727

359

0.65

[246

]

hot

gas

desu

lphu

riza

tion

wit

hZn

OH

DS-

10.

002k

g su

lphu

r/s

USD

2000

130

0.65

[297

]

Com

men

ts:

PW

S-1:

pres

suri

zed

wat

ersc

rubb

ing

unit

for

biog

as,i

nclu

ding

offga

str

eatm

ent,α

deri

ved

from

data

for

500

m3/h

and

1000

m3/h

plan

tsba

sed

on[2

32]

AG

R-1

:do

uble

stag

eSe

lexo

luni

tfo

rco

al-fi

red

IGC

C,2

9237

kmol

/hga

s,37

.7%

CO

2,0

.6%

H2S,α

from

onlin

esu

pple

men

tary

mat

eria

lpro

vide

dfo

r[3

88]

AG

R-2

:Se

lexo

luni

tfo

rco

al-fi

red

IGC

C,3

6998

kmol

/hga

s,37

.0%

CO

2,0

.2%

H2S

AG

R-3

:Se

lexo

luni

tfo

rbi

omas

sga

sific

atio

n

SRU

-1:

Cla

us/S

cot

plan

t

HD

S-1:

ZnO

guar

dbe

dfo

rF

TS

from

biom

ass

gasi

ficat

ion

wit

h69

.5M

WL

HV

,ass

umin

gbi

omas

ssu

lphu

rco

nten

tfr

omsi

mul

atio

nF

B-w

ood-

1

231


Tabl

eA

.41:

Cos

tda

tafo

rw

ater

gas

shift

and

met

hane

stea

mre

form

ing

reac

tors

.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

sour

shift

reac

tor

WG

S-1

5.25

kmol

gas/

s,

1.51

kmol

CO

/s

€20

10

(USD

2007

)

3908

0.6

[297

][2

46]

clea

nsh

iftre

acto

rW

GS-

20.

667

kmol

gas/

s,

0.12

0km

olC

O/s

€20

10

(USD

2007

)

3908

0.6

[297

][2

97]

met

hane

stea

mre

form

erM

SR-1

1.30

2kg

H2/s

USD

2000

1652

00.

7[3

95]

met

hane

stea

mre

form

erM

SR-2

0.90

9kg

H2/s

USD

2001

1710

00.

7[3

96]

Com

men

ts:

WG

S-1:

sour

shift

,inc

ludi

nghe

atex

chan

ger,

3778

4km

ol/h

raw

gas

wit

h28

.7%

CO

,2un

its,

11.3

MU

SD20

07.

Thi

sre

sult

sin

4.70

M€

2010

for

a5.

25km

ol/s

unit

,of

whi

chth

eco

stfo

rth

esy

ngas

cool

ers

(16

MW

)is

subt

ract

ed.

WG

S-2:

clea

nsh

ift,2

400

kmol

/hsy

ngas

.18

%C

Oas

sum

ed(b

ased

onsi

mul

atio

nF

B-w

ood-

1)

MSR

-1:

met

hane

stea

mre

form

ing

unit

,co

mpr

isin

gst

eam

refo

rmer

,w

ater

gas

shift

reac

tor

and

hydr

ogen

puri

ficat

ion,

H2

prod

ucti

on6.

3t/

h,op

erat

ing

pres

sure

30ba

r,41

.3M

USD

.Bas

edon

[396

,p.

32],

itis

assu

med

that

the

refo

rmer

cont

ribu

tes

40%

ofth

eto

talc

ost

for

refo

rmer

,shi

ftan

dpu

rific

atio

n.It

isas

sum

edth

at75

%of

the

H2

ispr

oduc

edin

the

refo

rmer

and

25%

inth

esh

iftre

acto

r.

MSR

-2:

met

hane

stea

mre

form

erfo

rsy

ngas

from

Bat

elle

biom

ass

gasi

fier

wit

h68

.75

t/h

dry

biom

ass

feed

.A

ssum

ing

aC

H4

and

othe

rlig

hthy

droc

arbo

nsyi

eld

of9.

911

kmol

/tfe

ed,d

ryfo

rth

eB

atel

lega

sifie

rba

sed

on[2

97].

Ass

umin

ga

conv

ersi

oneffi

cien

cyof

80%

for

the

refo

rmer

,thi

sle

ads

toa

H2

prod

ucti

onof

47.5

7/t

feed

,dry

.

232


Tabl

eA

.42:

Cos

tda

tafo

rga

stu

rbin

esy

stem

s.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

larg

e-sc

ale

[397

]

GT

air

com

pres

sor

GA

C-1

241

MW

shaf

t€

2010

(USD

2006

)

8046

1.1

0.89

(PE

C)

GT

com

bust

ion

cham

ber

GC

C-1

809.

7M

WH

HV

€20

10

(USD

2006

)

1455

1.1

0.89

(PE

C)

GT

expa

nder

GE

X-1

528

MW

shaf

t€

2010

(USD

2006

)

1422

81.

10.

89(P

EC

)

GT

gene

rato

r,el

ectr

ic

inst

alla

tion

and

auxi

liari

es

GG

E-1

284.

7M

Wel

€20

10

(USD

2006

)

2177

31.

10.

89(P

EC

)

med

ium

-sca

le[3

97]

GT

air

com

pres

sor

GA

C-2

35.8

MW

shaf

t€

2010

(USD

2006

)

1664

1.1

0.89

(PE

C)

GT

com

bust

ion

cham

ber

GC

C-2

97.4

MW

HH

V€

2010

(USD

2006

)

301

1.1

0.89

(PE

C)

GT

expa

nder

GE

X-2

67.2

MW

shaf

t€

2010

(USD

2006

)

2942

1.1

0.89

(PE

C)

GT

gene

rato

r,el

ectr

ic

inst

alla

tion

and

auxi

liari

es

GG

E-2

30.9

MW

el€

2010

(USD

2006

)

4520

1.1

0.89

(PE

C)

Com

men

ts:

larg

e-sc

ale:

spec

ific

cost

for

aga

stu

rbin

esy

stem

incl

udin

gst

ack:

183

USD

/kW

,ba

sed

on3

unit

sw

ith

capa

citi

esof

270-

334

MW

el(M

701G

,SG

T5-

4000

F,

M70

1F).

Thi

sre

sult

sin

45.4

6M

€20

10fo

ra

gas

turb

ine

syst

emw

ith

284.

7M

Wel

.

med

ium

-sca

le:

spec

ific

cost

for

aga

stu

rbin

esy

stem

incl

udin

gst

ack:

348

USD

/kW

,ba

sed

on6

unit

sw

ith

capa

citi

esof

25-4

5M

Wel

(SG

T-8

00,

SGT

-700

,SG

T-6

00,

OG

T25

000,

RB

211-

6562

DL

E,P

G65

91(C

)).T

his

resu

lts

in9.

41M

€20

10fo

ra

gas

turb

ine

syst

emw

ith

30.9

MW

el.

The

tota

lco

stis

split

betw

een

the

gas

turb

ine

com

pone

nts

base

don

[398

,p.

180]

:co

mpr

esso

r17

.7%

,co

mbu

stio

nch

ambe

r3.

2%,

expa

nder

31.3

%,

gene

rato

r8.

6%,

elec

tric

inst

alla

tion

and

auxi

liari

es39

.3%

.

Xre

fof

the

indi

vidu

alco

mpo

nent

sar

eta

ken

from

asi

mul

atio

nof

ana

tura

lgas

fired

gas

turb

ine.

CB

M/P

EC

isba

sed

onda

tafr

om[3

09].α

isas

sum

edto

be0.

89,a

ltho

ugh

data

from

[397

]ind

icat

esth

atbe

twee

n15

0an

d33

0M

W,α

iscl

ose

to1.

0.

233


Tabl

eA

.43:

Cos

tda

tafo

rm

isce

llane

ous

pow

erpl

ant

equi

pmen

t.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

HR

SGdu

ctw

ork,

foun

dati

ons

HR

S-1a

524.

8M

Wth

USD

2007

4681

0.9

[246

]

HR

SGst

ack

HR

S-1b

3818

t gas

/hU

SD20

0759

050.

9[2

46]

HR

SGdu

cts,

casi

ng,s

tack

HR

S-2

[kg g

as/s

]U

SD19

85P

EC

(m)

2.0

[383

][3

83]

HR

SGdr

ums

HR

D-1

[kg s

team

/s]

USD

1985

PE

C(m

)2.

0[3

83]

[383

]

feed

wat

ersy

stem

FW

S-1

240.

8kg

/s€

2010

(USD

2007

)

2216

40.

8[2

46]

cool

ing

wat

ersy

stem

CW

S-1

562

MW

thU

SD20

0730

067

0.7

[246

]

acce

ssor

yel

ectr

icpl

ant

AE

P-1

497

MW

elU

SD20

0710

2070

0.6

[246

]

Com

men

ts:

HR

S-1:

=H

RS-

1a+

HR

S-1b

HR

S-2:

cost

func

tion

PE

C=

0.52

0m1.

2.

Val

idat

ion

ofco

stfu

ncti

onse

eH

X-3

inT

able

A.2

7

HR

D-1

:co

stfu

ncti

onP

EC

=10

.957

m.

Val

idat

ion

ofco

stfu

ncti

onse

eH

X-3

inT

able

A.2

7

FW

S-1:

incl

udin

gfe

edw

ater

pum

psan

dde

aera

tor,

28.9

MU

SD20

07,o

fw

hich

the

feed

wat

erpu

mps

(2.7

M€

2010

)ar

esu

btra

cted

CW

S-1:

cool

ing

tow

er,c

oolin

gw

ater

pum

ps

AE

P-1

:tr

ansf

orm

ers,

swit

chge

ar,e

mer

genc

ydi

esel

gene

rato

r

234


Tabl

eA

.44:

Cos

tda

tafo

rau

xilia

ryun

its:

air

sepa

rati

on,w

aste

wat

ertr

eatm

ent,

engi

ne,r

efri

gera

tion

.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αso

urce

s

ASU

cold

box,

inte

grat

edA

SU-1

21.2

8kg

O2/s

USD

2003

2270

0(P

EC

)0.

5[3

88]

ASU

cold

box,

not

inte

grat

edA

SU-2

21.2

8kg

O2/s

USD

2003

2800

0(P

EC

)0.

5[3

88]

aero

bic

trea

tmen

tof

HT

Cw

aste

wat

erW

WT

-10.

001

m3/s

USD

2002

200

0.58

9[3

08,F

ig.

B-9

]

CH

Pen

gine

mod

ule

EN

G-1

100-

2000

kWel

€20

07f(

W)

[232

,p.

90]

refr

iger

atio

nm

achi

neF

RI-

120

kWco

olin

gU

SD20

0419

0.92

6[3

05,F

ig.

5.42

]

Com

men

ts:

ASU

-1,

ASU

-2:

cryo

geni

cA

SUco

ldbo

x,76

.6t/

hkg

/spu

reO

2.

ASU

-1is

oper

ated

atel

evat

edpr

essu

re.

Val

ues

give

nfo

rP

EC

isin

stal

led

cost

(IC

),C

BM

/PE

C=

1.2

(IC

+20

%fo

ren

gine

erin

g).

PE

Can

dα

from

supp

lem

enta

rym

ater

ialo

f[3

88]

WW

T-1

:sm

allp

acka

ged

was

tew

ater

trea

tmen

tpl

ant

incl

udin

gsc

reen

ing,

aera

tion

,cla

rific

atio

n,ch

lori

nati

on,a

erob

icsl

udge

dige

stio

n.α

deri

ved

for

2.0

E-5

to1.

0E

-3m

3/s

from

[308

,Fig

.B

-9]

EN

G-1

:co

stfu

ncti

on:

c=

4538

W(−

0.33

)e

lw

here

cis

spec

ific

CB

Min

[€20

07/k

Wel

],an

dW

elis

the

elec

tric

alpo

wer

capa

city

in[k

W],

deri

ved

from

data

for

5pl

ants

base

don

[232

,p.

90]

FR

I-1:

air-

cool

edm

echa

nica

lref

rige

rati

onun

it,w

itho

utev

apor

ator

,ref

rige

rant

tem

pera

ture

-5°C

,α

deri

ved

for

5-10

0kW

from

[305

,Fig

.5.

42]

Tabl

eA

.45:

Cos

tda

tafo

rIG

CC

offsi

teco

st.

type

code

Xre

fcu

rren

cy,y

ear

Cre

f[1

03]

CB

M/P

EC

αf M

sour

ces

site

deve

lopm

ent

OSC

-115

91M

WH

HV

USD

2007

1491

20.

65[3

27]

[246

]

anci

llary

build

ings

OSC

-215

91M

WH

HV

USD

2007

1499

10.

65[3

27]

[246

]

land

OSC

-315

91M

WH

HV

USD

2007

900

0.65

[327

][2

46]

235


A.3.3 Offsite costs

The offsite cost of the upgrading plants are estimated as percentages of the purchasedequipment cost of a defined reference case for each technology, and scaled to other plantcapacities with Equation 3.11 and the scaling exponents given in Table A.46. The resultingoffsite costs of the references cases are given in Table A.47. PEC is estimated based onthe CBM.

For the IGCC plants, utilities are included in the CBM (feedwater and cooling watersystems, accessory electric plant, feedstock and ash handling and storage). Cost functionsfor land, ancillary buildings and site development are given in Table A.45.

Table A.46: Scaling exponents α for offsite costs, based on [327], and offsite cost as % ofPEC in the reference cases of the upgrading plants.

Offsite cost α % of PECland 0.7 10%ancillary buildings 0.65 20%site development 0.65 5%utilities 0.65 10%

Table A.47: Reference cases for the calculation of the offsite cost.

WP, TOR AD HTC b) CHP c)

reference case WP-1.0-m ADP-3.0-s HTC-1.00-m CHP-1capacity [MWHHV,biomass] 64.8 9.51 54.3 38.70CBM [M€] 4.81 4.72 20.25 10.63CBM/PEC a) [-] 2.46 2.46 2.88 2.88offsite cost [M€] 0.88 0.86 3.40 1.66

a) Based on [306].b) For HTC-5.00 , land cost are assumed to be zero because the land otherwise used for EFB dumping sitescan be used for the HTC plant. Ancillary buildings are 10% PEC.c) For the integrated HTC/CHP plants CHPB-3.1 to CHPB-3 .3, it is assumed that the offsite costs equalthose of the standalone CHP plant plus 30% of the offsite cost of the standalone HTC plant HTC-3.00-s.

236


A.3.4 Costs of auxiliary energy and consumables

Table A.48: Cost of auxiliary energy and other consumables.

required for sources and comments

electricity, Germany upgrading [€/MWh] 100.00 for industrial consumers

electricity, Malaysia HTC-5.00 [€/MWh] 82.98 medium voltage general industrial

tariff [399]

electricity, feed-in from

biomass-fired CHP

ADP, WP-1.2 ,

WP-1.3,

CHPB

[€/MWh] 80.00 1)

oil palm shells HTC-5.00 [€/tFM] 20.00 [32]

iron oxide for AD ADM, ADP [€/m3STP ] 0.0046 [232, p. 75]

digestate spreading ADM, ADP [€/m3] 0.00 2)

biomethane feed-in ADM, ADP [€/MWhHHV] 1.452 [232, p. 87]

biomass ash disposal at

power station

WP, TOR,

HTC, ADP

[€/t] 50.00 3)

chemicals, water, waste

water

WP, TOR,

HTC, ADP

[€/tbiofuel] 0.500 waste water (after treatment) and ash

disposal, conditioner for pelletizing

ZnO for hot gas

desulphurization

FB-IGCC [€/kgS] 0.124

chemicals, water, waste

water

IGCC, PC [€/tbiofuel] 0.500 waste water and ash disposal, solvent

for AGR1) 78–83 €/MWh base tariff according to German Renewable Energy Sources Act (2009) for biomass-fired CHPplants with a capacity of 0.5-20 MWel.2) may be positive or negative. Revenues of 6 €/m3 from the sale of the digestate [218] are reported as well as netcost of 4 €/m3 for spreading [102, p. 197].3) may range from 20-500 €/t, depending on application or disposal [270, 293, 394, 400], [401, p. 348].

A.3.5 Operating labour

Labour rates for Germany and Malaysia and scaling exponents for the respective types oflabour are given in Table A.49.

Additional labour in the preparation yard is assumed to be required for sizing and sortingfor plants processing PGW and MOW.

The estimated labour requirements for the reference plants of each technology are givenin Table A.50 and Table A.51.

The annual labour hours are calculated with the required number of persons per shift andan assumed working time of 8520 h/a, comprising the 7008 h/a full-load operating hoursof the plants plus an allowance for repairs and revisions.

A minimum of one operator per shift is assumed independently of the plant capacity for theplants producing solid biofuels to ensure continuous supervision. For the plants processingwaste and grass, a minimum of one worker per shift is assumed for screening and sorting.

237


Table A.49: Hourly cost of labour and scaling exponents for labour requirement. Labourrates for Germany are based on data from the Federal Statistical Office of Germany, andfor Malaysia on [402, 403] for technicians and on [404] for unskilled workers. Scalingexponents are based on [315, p. 80].

Germany Malaysia α

[€/h] [€/h]plant operator 27.625 2.123 0.3worker for biomass yard, feedstock sizing and sorting 20.750 0.512 0.5

Table A.50: Requirement of plant operators per shift.

Xref nref source and comments

[MWfeedstock,HHV] [persons /shift]

upgrading plants

WP 10.86 1.0 [313],[143, p. 144]

TOR 10.86 1.0 like WP

ADM 10.86 0.3 estimated, based on data from [232]

ADP 10.86 1.0 like WP

HTC 10.86 3.0 based on estimate for fast pyrolysis plant

[315] and WP

WP with CHP (WP-1.2,-1.3 ) 100 2.9 based on data for CHP plant [314] and WP

wood-fired CHP 38.7 1.0 [314]

HTC with CHP (CHPB) 80% of standalone HTC and CHP plant

power plants

IGCC with CCS 1616 16.0 [316]

IGCC without CCS 1547 15.0 [316]

PC without CCS 1547 15.0

CCGT with CCS 1103 6.3 [316]

CCGT without CCS 1102 5.0 [316]

Table A.51: Requirement of workers for biomass yard, product handling, feedstock sizingand sorting, per shift.

Xref nref source and comments

[MWfuel,HHV] [persons/shift]

CHP and upgrading plants with solid products 10.86 1.0 [314]

ADM 10.86 0.5 no product handling required

additional labour for waste and grass 10.86 1.0

IGCC processing raw wood chips 10.86 1.0 like upgrading plants

IGCC processing pellets 0.0

238

A.4 Exergetic efficiency definitions and auxiliary costing equations

A.3.6 Levelized product costs

Table A.52: Escalation rates.

natural gas, diesel, electricity [% p.a] 0.5%biomass, coal [% p.a] 0.3%remuneration for electricity from CHP plant 1) [% p.a] 0.0%other [% p.a] 0.0%

1) The escalation rate for electricity feed-in from biomass-fired CHP plants is assumed to be zero, repres-enting a constant feed-in tariff over the economic plant life.

A.4 Exergetic efficiency definitions and auxiliary costingequations

A.4.1 Upgrading plants

The indices in Table A.54 refer to the flow stream numbers from Figure 4.8, those inTable A.53 to Figure 4.3 (torrefaction), Figure 4.4 (ADM) and Figure 4.5 (ADP).

For the anaerobic digester, the efficiency εII according to equation Equation 3.28 cannotbe defined because the specific chemical exergy of the biogas is lower than that of thefeedstock biomass, due to the high CO2 content of the biogas. However, εII can be definedfor the overall plant, where the biogas is upgraded to biomethane.

Table A.53: Exergetic efficiency definitions for torrefaction and anaerobic digestion equip-ment.

component exergetic efficiency

torrefaction reactorεreactor =

ms,13(eCHs,13−eCH

s,3 )(ms,3−ms,13)eCH

s,3 +(

ECHlg,3−ECH

lg,13

)+(EP H

3 −EP H13 )+(E24−E14)

anaerobic digestorεdigester = E8

E2−E7+W5+(E36−E37)

A.4.2 BECCS plants

For the syngas production plants discussed in section 5.1, the exergetic efficiency is definedas

ε =Eclean gas + EΔQ,net + ECO2

Efeed + Eair + Esteam + Ewater + Esand + Edolomite + Wel

(A.1)

239


where subscript ΔQ, net relates to the thermal energy in the form of steam and hot waterwhich is generated in the raw gas coolers, plus the thermal energy of the reformer exhaustgas (stream 242 in Figure 5.2) minus the thermal energy consumed by air separation unit,acid gas removal and drier. Esteam and Ewater refer to the steam used as gasification agentand the water for the scrubber. Sand and dolomite are the makeup for the fluidized bedmaterial lost with the ash.

240

A.4 Exergetic efficiency definitions and auxiliary costing equations

Tabl

eA

.54:

Exe

rget

iceffi

cien

cyde

finit

ions

and

auxi

liary

cost

ing

equa

tion

sfo

rH

TC

-1.0

0an

dH

TC

-3.0

0.

exer

geti

ceffi

cien

cyau

xilia

ryco

stin

geq

uati

ons

HT

Cre

acto

rε K

14=

ms

,42(e

CH

s,4

2−

eC

Hs

,10)+

(EP

H42

+E

PH

80−

EP

H10

)(m

s,1

0−

ms

,42)e

CH

s,1

0+

( EC

Hlg

,10

−E

CH

lg,4

2−

EC

Hlg

,80

)cC

Hlg

,80

=cC

Hlg

,10

CC

Hs

,80

=0

CC

Hlg

,42

=m

lg,4

2( eC

Hlg

,10cC

Hlg

,10

+( eC

Hlg

,42

−eC

Hlg

,10) cC

Hs

,10)

CP

H42

=m

lg,4

2( eP

H10

cPH

10+

( ePH

42−

ePH

10) cC

Hs

,10)

CC

Hlg

,80

=m

lg,4

2( eC

Hlg

,10cC

Hlg

,10

+( eC

Hlg

,80

−eC

Hlg

,10) cC

Hs

,10)

flash

tank

ε K17

=m

15(e

PH

15−

eP

H6

)m

16(e

PH

6−

eP

H16

)+( E

CH

lg,6

−E

CH

lg,1

5−

EC

Hlg

,16

)cP

H16

=cP

H6

cCH

lg,1

5=

cCH

lg,6

cCH

lg,1

6=

cCH

lg,6

CC

Hs

,15

=0

CC

Hs

,16

=C

CH

s,6

slur

ry/s

team

mix

erε K

5=

m76

(eP

H19

−e

PH

76)

m47

(eP

H47

−e

PH

19)+

( EC

Hlg

,76

−E

CH

lg,4

7−

EC

Hlg

,19

)cC

Hlg

,19

=C

CH

lg,7

6+

CC

Hlg

,47

EC

Hlg

,76

+E

CH

lg,4

7C

CH

s,1

9=

CC

Hs

,76

filte

rpr

ess

not

defin

ed

cCH

lg,5

3=

cCH

lg,9

6

cCH

lg,5

4=

cCH

lg,9

6

cPH

53=

cPH

96

cPH

54=

cPH

96

CC

Hs

,54

=0

drie

rno

tde

fined

cCH

lg,6

8=

CC

Hlg

,64

+C

CH

lg,5

3E

CH

lg,6

4+

EC

Hlg

,53

cCH

lg,6

8=

CP

H64

+C

PH

53E

PH

64+

EP

H53

CC

Hs

,68

=0

CC

Hs

,99

=0

cPH

60=

cPH

53

cCH

lg,6

0=

cCH

lg,5

3

cPH

99=

cPH

53

cCH

lg,9

9=

cCH

lg,5

3

241

B Biomass upgrading plant data

B.1 Wood pelletizing plant models

B.1.1 Cost data

WP-1.0-l, WP-1.2-l and WP-1.3-l employ 5 pellet presses with a capacity of 13.3 t/h eachdue to the limited maximum capacity per unit for which pellet presses are commerciallyavailable. For the same reason, WP-1.2-l and WP-1.3-l employ 2 belt driers, including heatexchanger and fan. In addition to the equipment listed in Table B.2 and Table B.3, WP-1.2 employs an SSD drier (cost function SSD-1 ), a stainless-steel steam compressor (costfunction CMP-1 ), and a small wood boiler for process steam production (cost functionWB-1 ). The steam compressor is made from stainless steel because the steam containsevaporated organic compounds from wood.

Table B.1: Investment costs for wood pellet plant models using SR wood.

1.0-s 1.0-m 1.0-l 1.1-s 1.1-m 1.2-m 1.2-l 1.3-m 1.3-l

pellet production [kt/a] 15.6 77.9 389.3 15.6 77.9 77.9 389.3 77.9 389.3

wood consumption [kt/a] 33.5 167.3 836.3 29.1 145.3 191.7 958.5 182.3 911.5

drier [M€] 0.71 1.82 4.64 1.75 4.48 2.94 10.03 2.94 10.03

wood furnace [M€] 0.16 0.47 1.36

drier HX [M€] 0.33 1.73 0.33 1.73

drier air fan [M€] 0.00 0.01 0.09 0.06 0.48 0.06 0.48

fluidization st. comp. [M€] 0.08 0.29

heating steam comp. [M€] 0.65 2.17

hammer mill [M€] 0.12 0.37 0.37 0.12 0.37 0.37 0.37 0.37 0.37

pellet press [M€] 0.49 1.45 7.25 0.49 1.45 1.45 7.25 1.45 7.25

pellets stor. & handl. [M€] 0.24 0.69 0.69 0.24 0.69 0.69 0.69 0.69 0.69

wood boiler [M€] 0.03 0.11 4.74 16.66 4.66 16.37

steam turbine [M€] 1.19 3.38 1.19 3.38

feedwater pump [M€] 0.01 0.03 0.15 0.48 0.15 0.48

total CBM [M€] 1.72 4.81 14.39 3.37 9.57 11.93 41.07 11.85 40.78

offsite cost [M€] 0.30 0.88 2.55 0.28 0.80 1.86 5.40 1.80 5.23

fees & contingencies [M€] 0.26 0.74 2.16 0.50 1.44 1.79 6.16 1.78 6.12

start-up [M€] 0.14 0.46 1.82 0.18 0.61 0.63 2.35 0.61 2.29

working capital [M€] 0.55 2.33 11.15 0.53 2.26 2.54 12.00 2.43 11.43

AFUDC [M€] 0.23 0.65 1.91 0.41 1.18 1.56 5.26 1.54 5.21

residual value (NPV) [M€] -0.12 -0.53 -2.58 -0.09 -0.44 -0.52 -2.55 -0.49 -2.41

TCI [M€] 3.08 9.34 31.39 5.19 15.43 19.79 69.70 19.52 68.64

243

Chapter B Biomass upgrading plant data

Tabl

eB

.2:W

P-1.

0-s

equi

pmen

tlis

tw

ith

inve

stm

ent

cost

s.

com

pone

ntun

itX

sim

f dn

Xun

itC

BM

unit

[k€]

CB

M[k

€]sp

ecifi

cati

ons

cost

func

tion

drum

drie

r[k

g ev/h

]17

7812

0%1

2133

711

711

RD

D-1

=65

9k€

,R

DD

-2=

547

k€,

RD

D-3

=92

6k€

woo

dfu

rnac

e[k

W]

2100

120%

121

0016

316

3W

F-1

drie

rai

rfa

n[m

3/s

]6.

9711

0%1

7.67

44

AF

-1

ham

mer

mill

[kg/

h]22

2212

0%1

2667

120

120

HM

-1

pelle

tpr

ess

[kg/

h]22

2212

0%1

2667

489

489

incl

udin

gpe

llet

cool

er,

siev

ing

and

scre

enin

g

x(P

P-1

,PP

-2)+

PC

-1+

PSS

-1

PP

-1=

319

k€,

PP

-2=

313

k€

PC

-1=

147

k€,

PSS

-1=

26k€

pelle

tsst

orag

e&

hand

ling

[kg/

h]22

2212

0%1

2667

238

238

PS-

1

tota

lC

BM

1724

Tabl

eB

.3:W

P-1.

2-m

equi

pmen

tlis

tw

ith

inve

stm

ent

cost

s.

com

pone

ntun

itX

sim

f dn

Xun

itC

BM

unit

[k€

]C

BM

[k€

]sp

ecifi

cati

ons

cost

func

tion

belt

drie

r[k

g ev/h

]88

8912

0%1

1066

729

4129

41B

D-1

=43

48k€

,BD

-2=

1533

k€

drie

rH

X[m

2]

1352

115%

115

5533

433

4C

S/C

SH

X-1

drie

rai

rfa

n[m

3/s

]11

8.33

110%

113

0.17

6262

CS

AF

-1

ham

mer

mill

[kg/

h]11

111

120%

113

333

370

370

HM

-1

pelle

tpr

ess

[kg/

h]11

111

120%

113

333

1450

1450

incl

udin

gpe

llet

cool

er,

siev

ing

and

scre

enin

g

x(P

P-1

,PP

-2)+

PC

-1+

PSS

-1

pelle

tsst

orag

e&

hand

ling

[kg/

h]11

111

120%

113

333

688

688

PS-

1

boile

r[M

W]

2012

0%1

2447

4047

4088

bar,

500°

CW

B-2

stea

mtu

rbin

e[M

W]

3.1

110%

13.

411

9511

95ST

B-1

=13

55k€

,ST

B-2

=10

34

k€

feed

wat

erpu

mp

[kW

]61

.922

0%2

68.0

7515

0P

-1

tota

lC

BM

1192

9

244

B.2 Torrefaction plant model


B.2.1 Simulation data from TOR-1.0

Table B.4: Composition (d.b.) of the torrefied wood for the torrefaction simulation model,and measured data from Prins et al. [141].

raw wood torrefied wood

simulation [141] simulation [141]

c 47.49% 47.20% 51.62% 51.30%

h 6.42% 6.10% 6.28% 5.90%

n 0.07% 0.30% 0.08% 0.40%

s 0.05% 0.06%

o 44.18% 45.10% 39.92% 40.90%

ash 1.78% 1.30% 2.04% 1.50%

Table B.5: TOR-1.0 flow stream data.

stream no. 1 3 5 6 7 8 13 14 15 16 17 18 19 20T [°C] 15.0 80.0 110.1 110.1 110.1 18.8 250.0 250.0 269.1 269.1 269.1 40.0 1042.0 600.0p [bar] 1.013 1.013 1.013 1.013 1.013 1.05 1.013 1.013 1.153 1.153 1.153 1.013 1.05 1.05m [kg/h] 4000.00 2500.00 10696.12 6333.11 4363.01 3415.75 1799.22 3220.78 2520.00 700.78 3220.78 1799.22 4833.13 9196.14H [kW] -11694.00 -4969.80 -20054.00 -11874.00 -8180.20 -86.57 -2269.40 -10618.00 -8282.90 -2303.40 -10586.00 -2447.00 -4593.30 -12766.00O2 [kg/h] 0.00 0.00 381.52 225.90 155.62 790.63 0.00 0.00 0.00 0.00 0.00 0.00 225.89 381.52CO2 [kg/h] 0.00 0.00 1341.23 794.14 547.10 0.00 0.00 265.28 207.56 57.72 265.28 0.00 794.13 1341.23CO [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 27.48 21.50 5.98 27.48 0.00 0.00 0.00H2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 2000.00 500.00 4575.16 2708.93 1866.24 21.62 53.98 2571.28 2011.82 559.46 2571.28 53.98 1208.99 3075.23CH4 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/h] 0.00 0.00 4397.56 2603.77 1793.79 2603.50 0.00 0.00 0.00 0.00 0.00 0.00 2603.74 4397.53H2S [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00C2H4O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CH4O [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00torr. wood DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 1745.24 0.00 0.00 0.00 0.00 1745.24 0.00 0.00biomass DM [kg/h] 2000.00 2000.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.000 0.010 0.536 0.317 0.219 0.003 0.055 0.491 0.406 0.113 0.519 0.001 1.227 1.260ECH

s [MW] 11.409 11.409 0.000 0.000 0.000 0.000 10.769 0.000 0.000 0.000 0.000 10.769 0.000 0.000ECH

lg [MW] 0.028 0.007 0.181 0.107 0.074 0.004 0.001 1.766 1.382 0.384 1.766 0.001 0.089 0.163

stream no. 21 24 25 26 27 29 [kWel]T [°C] 500.8 540.0 15.0 800.0 114.5 15.0 W1 36.9 torr. gas compressorp [bar] 1.05 1.133 1.013 1.013 1.05 1.013 W2 9.3 exhaust gas compressorm [kg/h] 9196.14 2520.00 728.29 11.69 4363.01 3415.75 W3 4.5 air fanH [kW] -13141.00 -7908.20 -2129.20 3.19 -8172.70 -90.54 W4 90.0 pellet press O2 [kg/h] 381.52 0.00 0.00 0.00 155.62 790.63 W5 4.8 cooler fanCO2 [kg/h] 1341.23 207.56 0.00 0.00 547.10 0.00 W6 145.4 total power consumptionCO [kg/h] 0.00 21.50 0.00 0.00 0.00 0.00H2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 3075.23 2011.82 364.14 0.00 1866.24 21.62 heat losses [kWth]CH4 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 torrefaction reactor 12.0N2 [kg/h] 4397.53 0.00 0.00 0.00 1793.79 2603.50 furnace 71.0H2S [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 drier 189.0C2H4O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 cooler duty [kWth]CH4O [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 torr. wood cooler 177.6torr. wood DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00biomass DM [kg/h] 0.00 0.00 364.14 0.00 0.00 0.00ash [kg/h] 0.00 0.00 0.00 6.50 0.00 0.00EPH [MW] 1.016 0.618 0.000 0.002 0.225 0.000ECH

s [MW] 0.000 0.000 2.077 0.000 0.000 0.000ECH

lg [MW] 0.163 1.382 0.005 0.049 0.074 0.004

245


B.2.2 Cost data

The maximum capacity of a torrefaction reactor is around 50–60 kt/a [144]. It is assumedto be 10 t/h of dried biomass for the investment cost estimate, resulting in 2 reactors with7.5 t/h each in TOR-1.0-m and 8 reactors with 9.38 t/h each in TOR-1.0-l. In TOR-1.0-l,there are 4 pellet presses with a processing capacity of 13.5 t/h each.

246


Tabl

eB

.6:T

OR

-1.0

-meq

uipm

ent

list

wit

hin

vest

men

tco

sts.

com

pone

ntun

itX

sim

f dn

Xun

itC

BM

unit

[k€

]

CB

M

[k€

]

spec

ifica

tion

sco

stfu

ncti

on

torr

efac

tion

reac

tor

[kg/

h]12

500

120%

275

0012

5225

05T

OR

-1

furn

ace

[kW

HH

V]

1181

312

0%1

1417

561

161

1w

ood

furn

ace

wit

had

diti

onal

burn

erfo

rto

rref

acti

onga

s

WF

-1,p

lus

20%

for

addi

tion

alga

sbu

rner

heat

exch

ange

r[m

2]

738

115%

184

835

335

3SS

/CS

HX

-1

prod

uct

cool

er[m

2]

16.5

115%

119

129

129

SSA

CO

-1

torr

efac

tion

gas

com

pres

sor

[kW

shaf

t]15

611

0%1

176

432

432

SSC

MP

-1

mot

or[k

Wel

]18

511

0%1

203

8686

EM

-1

exha

ust

gas

com

pres

sor

[kW

shaf

t]38

110%

142

129

129

SSC

MP

-1

mot

or[k

Wel

]45

110%

150

3737

EM

-1

drie

rai

rfa

n[m

3/s

]3.

911

0%1

4.3

33

AF

-1

drum

drie

r[k

g ev/h

]75

0012

0%1

9000

1644

1644

RD

D-1

=15

25k€

,

RD

D-2

=12

66k€

,

RD

D-3

=21

42k€

pelle

tpr

ess

[kg/

h]89

9612

0%1

1079

512

5712

57in

clud

ing

pelle

tco

oler

,sie

ving

and

scre

enin

g

x(P

P-1

,PP

-2)+

PC

-1+

PSS

-1

pelle

tsst

orag

e&

hand

ling

[kg/

h]89

9612

0%1

1079

559

959

9P

S-1

tota

lC

BM

7785

247


B.3 Anaerobic digestion plant models

B.3.1 Model for the separation of press fluid and press cake for ADP-3.0

The model is based on experimental data from [234]. The data comprises mass flows oforganic matter, ash, N and S into press fluid and press cake in the mechanical separationstep. Data on the composition includes dry matter, ash, N and S content of feedstock,press fluid and press cake. No data is given on the C, H and O content of the products.

Table B.7 shows the mass flows into the press fluid and press cake and the compositionof the products. For the data from [234], an O/C and H/C ratio from ryegrass [28] isassumed. The O/C and H/C ratios are assumed to be equal in feedstock, press fluid andpress cake.

B.3.2 Digestion model

The anaerobic digestion model for simulation cases ADP-3.0 and ADM-3.0 is based onexperimental data from [234]. For ADM-3.1, a high methane yield is assumed based onliterature data for grass from intensive cropping systems (see Table 2.10).

Since the data from [234] is not sufficient to form elemental balances for C, H and O, thefollowing procedure is employed to build a simple black box model.

The conversion factor γ∗i is defined, where γi,max is the maximum theoretical yield of

CH4, CO2 or H2O for the complete conversion of the feedstock according to equationEquation 2.4:

γ∗i =

γi,real

γi,max(B.1)

γ∗CH4 is calculated for data from [234].

Hydrolysis and acetogenesis consume water, while methanogenesis and acidogenesis formwater [227]. Since no data on the net water consumption was available, it is assumed that

γ∗H2O = γ∗

CH4 (B.2)

γ∗CO2 is then fitted to meet the measured conversion of organic dry matter from [234]. The

resulting biogas composition of 46–52% CH4 and 48–54% CO2 seems plausible.

The resulting conversion factors γ∗CH4 , γ∗

H2O and γ∗CO2 are then applied to the feedstock

and press fluid composition used for the simulations models in this work.

It is assumed that 30% of the substrate sulphur reacts to H2S.

The composition of the digestate is calculated by difference from the elemental balances.The heat of reaction is calculated by difference from the energy balance.

248


Table B.7: Mass flows into and composition of press fluid and press cake, for simulationmodel ADP-1, and based on data from literature [234].

literature data ADP-3.0

feed composition (d.b.)

c [w% ] 47.12% 42.54%

h [w% ] 5.72% 5.94%

n [w% ] 1.27% 1.14%

s [w% ] 0.13% 0.16%

o [w% ] 38.94% 44.62%

Cl [w% ] 0.40% 0.00%

ash [w% ] 6.41% 5.60%

HHV (d.b.) [MJ/kg] 19.03 17.12

mass flows into press fluid

dry matter [–] 27.70% 27.70%

ash [–] 43.50% 43.50%

Cl [–] 90.10% 90.10%

N [–] 38.70% 38.70%

S [–] 56.30% 56.30%

press cake composition (d.b.)

c [w% ] 48.14% 43.21%

h [w% ] 5.85% 6.03%

n [w% ] 1.08% 0.97%

s [w% ] 0.08% 0.10%

o [w% ] 39.79% 45.32%

Cl [w% ] 0.05% 0.00%

ash [w% ] 5.01% 4.38%

HHV (dry) [MJ/kg] 19.47 17.41

press fluid composition (d.b.)

c [w% ] 44.45% 40.80%

h [w% ] 5.40% 5.69%

n [w% ] 1.77% 1.59%

s [w% ] 0.26% 0.33%

o [w% ] 36.74% 42.79%

Cl [w% ] 1.30% 0.00%

ash [w% ] 10.07% 8.79%

HHV (dry) [MJ/kg] 17.870 16.352

energy yield (HHV)

press fluid [–] 26.0% 26.5%

press cake [–] 74.0% 73.5%

249


Table B.8: Data for the anaerobic digestion model for simulation cases ADP-3.0, ADM-3.0 and ADM-3.1, and data from literature [234].

AD of press fluid AD of whole plant

literature ADP-3.0 literature ADM-3.0 ADM-3.1

conversion factors

γ∗CH4 [–] 91.8% 91.8% 45.8% 45.8% 85.0%

γ∗CO2 [–] 87.7% 55.9% 60.0%

γ∗H2O [–] 91.8% 45.8% 85.0%

S → H2S [–] 30.0% 30.0% 30.0%

CH4 yield [l/kgoDM] 445.13 426.60 226.47 197.01 365.29

oDM conversion [–] 85.00% 86.70% 53.80% 53.37% 63.70%

biogas composition

CH4 [mol%] 52.18% 46.1% 59.6%

CO2 [mol%] 47.72% 53.8% 40.3%

H2S [mol%] 0.100% 0.083% 0.058%

digestate composition

c [–] 19.91% 42.23% 29.05%

h [–] 2.21% 6.47% 2.23%

n [–] 7.61% 2.30% 2.86%

s [–] 1.09% 0.23% 0.28%

o [–] 27.15% 37.49% 51.53%

ash [–] 42.03% 11.29% 14.05%

HHV (d.b.) [MJ/kgDM] 5.85 18.24 7.13

mass yield [–] 20.9% 49.62% 39.87%

energy yields from digester substrate

biogas [–] 89.1% 86.9% 44.3% 43.2% 80.1%

digestate [–] 7.5% 52.9% 16.6%

heat of reaction [–] 5.6% 3.9% 3.3%

energy yields from feed

biogas [–] 23.2% 23.0% 44.3% 43.2% 80.1%

press cake [–] 74.0% 73.5% — — —

digestate [–] 2.0% 52.9% 16.6%

heat of reaction [–] 1.5% 3.9% 3.3%

250


B.3.3 Simulation data from ADM-3.0 and ADP-3.0

Table B.9: ADM-3.0 flow stream data.

1 2 3 4 5 6 7 8 9 10 11 12 13 14T [°C] 15.0 27.1 5.0 5.0 105.4 5.0 39.0 39.0 37.0 37.0 60.0 160.2 60.0 61.0p [bar] 1.013 1.013 1.013 1.013 3.000 1.013 1.013 1.013 1.013 1.013 3.000 8.000 8.000 7.000m [kg/h] 6666.7 13333.7 1161.4 1118.6 1118.6 42.8 12172.3 1161.4 3307.9 8858.3 1118.6 1118.6 1118.6 290.3H [kW] -23529 -52818 -2662 -2473 -2435 -189 -50384 -2620 -11477 -38916 -2453 -2413 -2454 -408O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 0.00 845.87 845.80 845.80 0.06 5.91 845.87 0.00 0.00 845.80 845.80 845.80 21.15CO [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 4666.67 11333.7 48.42 5.66 5.66 42.76 11173.8 48.42 2315.55 8858.29 5.66 5.66 5.66 5.66CH4 [kg/h] 0.00 0.00 266.13 266.13 266.13 0.00 0.09 266.13 0.00 0.00 266.13 266.13 266.13 263.47N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/h] 0.00 0.00 1.00 1.00 1.00 0.00 0.02 1.00 0.00 0.00 1.00 1.00 1.00 0.00SO2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00R134a [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00biomass DM [kg/h] 2000.00 2000.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00digestate DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 992.38 0.00 992.38 0.00 0.00 0.00 0.00 0.00ash [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.000 0.004 0.000 0.000 0.031 0.000 0.013 0.002 0.003 0.008 0.027 0.060 0.050 0.023ECH

s [MW] 10.062 10.062 0.000 0.000 0.000 0.000 5.252 0.000 5.252 0.000 0.000 0.000 0.000 0.000ECH

lg [MW] 0.065 0.157 3.928 3.928 3.928 0.001 0.156 3.928 0.032 0.123 3.928 3.928 3.928 3.795

15 16 17 18 20 22 23 24 25 26 27 29 30 32T [°C] 3.0 60.0 3.0 3.0 15.0 39.4 -10.0 -10.0 29.3 27.3 94.5 -10.0 24.1 49.0p [bar] 7.000 8.000 7.000 7.000 16.000 10.000 2.000 2.000 2.000 2.000 10.000 2.000 2.000 4.000m [kg/h] 290.3 828.3 285.0 5.3 285.0 1200.0 1200.0 900.0 900.0 1200.0 1200.0 300.0 300.0 2400.0H [kW] -422 -2046 -398 -24 -397 -2983 -2983 -2237 -2195 -2927 -2909 -746 -732 -10513O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 21.15 824.66 21.14 0.00 21.14 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 5.66 0.00 0.34 5.32 0.34 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 2400.00CH4 [kg/h] 263.47 2.66 263.47 0.00 263.47 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/h] 0.00 1.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00R134a [kg/h] 0.00 0.00 0.00 0.00 0.00 1200.00 1200.00 900.00 900.00 1200.00 1200.00 300.00 300.00 0.00biomass DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00digestate DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.022 0.026 0.022 0.000 0.031 0.012 0.010 0.007 0.004 0.005 0.019 0.002 0.001 0.005ECH

s [MW] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000ECH

lg [MW] 3.795 0.148 3.795 0.000 3.795 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.033

33 34 35 36 37 44 46 [kWel]T [°C] 64.8 49.0 59.3 62.6 49.0 37.0 37.0 W1 43.7 compressor (K5)p [bar] 4.000 4.000 4.000 4.000 4.000 1.013 1.013 W2 46.1 compressor (K7)m [kg/h] 2400.0 1600.0 1600.0 4000.0 4000.0 6667.0 2191.3 W3 16.2 compressor (K12)H [kW] -10471 -7008 -6991 -17462 -17521 -29289 -9627 W4 20.8 compressor (K18)O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 W5 123.2 stirrer (fermenter + conditioning)CO2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 W6 17.1 decanter (K13)CO [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 W7 113.4 pressurized water scrubber (pumps)H2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 W8 6.1 cooler fansH2O [kg/h] 2400.00 1600.00 1600.00 4000.00 4000.00 6667.00 2191.29 W9 386.5 totalCH4 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 heat losses [kWth]H2S [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 fermenter (K2) 80.4SO2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 refrigerant condenser (K19) 74.0R134a [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 decanter (K13) 23.8biomass DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00digestate DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.010 0.003 0.005 0.016 0.008 0.006 0.002ECH

s [MW] 0.000 0.000 0.000 0.000 0.000 0.000 0.000ECH

lg [MW] 0.033 0.022 0.022 0.056 0.056 0.093 0.030

251


Table B.10: ADP-3.0 flow stream data.

1 2 3 4 6 7 8 9 10 12 13 14 16T [°C] 15 59.95169 59.73431 37 48.42122 39 39 37 37 3 70 3 15p [bar] 1.013 1.013 1.013 1.013 1.013 1.013 1.013 1.013 1.013 1.013 1.013 1.013 1.013m [kg/h] 6666.667 36837.57 3057.074 30170.9 33780.49 33276.44 504.0496 386.3993 32890.04 483.9367 483.4323 20.11284 3652.979H [kW] -23529 -155060 -9136.8 -132520 -146340 -145700 -1107.6 -1307.4 -144460 -1038.8 -1027.2 -89.0186 -96.8331O2 [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 845.537CO2 [kg/h] 0.000 0.000 0.000 14.009 0.000 15.271 339.890 0.000 15.271 339.862 339.862 0.028 0.000CO [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000H2 [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000H2O [kg/h] 4666.7 34837.6 1611.1 30156.6 33226.5 33144.9 22.350 270.480 32874.4 2.266 2.266 20.085 23.126CH4 [kg/h] 0.000 0.000 0.000 0.269 0.000 0.293 141.305 0.000 0.293 141.305 141.305 0.000 0.000N2 [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 2784.32H2S [kg/h] 0.000 0.000 0.000 0.066 0.000 0.072 0.504 0.000 0.072 0.504 0.000 0.000 0.000SO2 [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000R134a [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000biomass DM [kg/h] 2000.00 2000.00 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000digestate DM [kg/h] 0.000 0.000 0.000 0.000 0.000 115.920 0.000 115.920 0.000 0.000 0.000 0.000 0.000press fluid oDM [kg/h] 0.000 0.000 0.000 0.000 554.004 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000press cake DM [kg/h] 0.000 0.000 1446.00 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000ash [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000EPH [MW] 0.000 0.122 0.008 0.028 0.066 0.036 0.001 0.000 0.031 0.000 0.001 0.000 0.000ECH

s [MW] 10.062 10.062 7.398 0.000 2.664 0.209 0.000 0.209 0.000 0.000 0.000 0.000 0.000ECH

lg [MW] 0.065 0.483 0.022 0.422 0.461 0.464 2.073 0.004 0.460 2.073 2.070 0.000 0.005

17 18 19 20 22 23 24 25 27 28 31 32 34T [°C] 454 697.0213 80 113.5593 40 40 15 1286.506 59.86268 90.10367 90.10367 64.40739 39.39086p [bar] 1.013 1.013 1.013 1.013 1.013 1.013 1.013 1.013 2 2.026 2.026 2.026 10m [kg/h] 4136.411 5731.954 1606.662 7182.365 1421.662 185 1419.988 1595.543 15000 15000 399 399 500H [kW] -2531.3 -2919.8 -2781.3 -9350.2 -2493.7 -324.5 -37.641 -388.55 -65528 -65021 -1729.6 -1741.1 -1243.1O2 [kg/h] 281.846 420.200 0.000 420.200 0.000 0.000 328.678 138.354 0.000 0.000 0.000 0.000 0.000CO2 [kg/h] 727.501 983.218 0.000 983.218 0.000 0.000 0.000 255.718 0.000 0.000 0.000 0.000 0.000CO [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000H2 [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000H2O [kg/h] 342.750 459.969 160.666 1910.4 142.166 18.500 8.990 117.219 15000.0 15000.0 399.000 399.000 0.000CH4 [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000N2 [kg/h] 2784.32 3868.25 0.000 3868.25 0.000 0.000 1082.3 1083.9 0.000 0.000 0.000 0.000 0.000H2S [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000SO2 [kg/h] 0.000 0.322 0.000 0.322 0.000 0.000 0.000 0.322 0.000 0.000 0.000 0.000 0.000R134a [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 500.000biomass DM [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000digestate DM [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000press fluid oDM [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000press cake DM [kg/h] 0.000 0.000 1446.00 0.000 1279.50 166.500 0.000 0.000 0.000 0.000 0.000 0.000 0.000ash [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000EPH [MW] 0.241 0.643 0.006 0.200 0.001 0.000 0.000 0.440 0.051 0.135 0.004 0.002 -0.002ECH

s [MW] 0.000 0.000 7.398 0.000 6.547 0.852 0.000 0.000 0.000 0.000 0.000 0.000 0.000ECH

lg [MW] 0.070 0.095 0.002 0.114 0.002 0.000 0.002 0.024 0.208 0.208 0.006 0.006 0.000

35 36 37 39 40 41 42 43 44 45 46 47 48T [°C] -6.90992 -1.55281 64.9408 40 64.80536 37 59.86299 90.10367 59.73431 59.73431 49.73431 74.67967 68p [bar] 2.26 2.26 10 1.013 1.013 1.013 2.026 2.026 2.026 1.013 1.013 1.013 1.013m [kg/h] 500 500 500 1606.662 30170.9 2719.144 15000 14601 14601 33780.49 30170.9 7182.365 30170.9H [kW] -1243.1 -1222.9 -1216.2 -2818.2 -131640 -11943 -65528 -63291 -63787 -145930 -132130 -9456.6 -131530O2 [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 420.200 0.000CO2 [kg/h] 0.000 0.000 0.000 0.000 0.000 1.263 0.000 0.000 0.000 0.000 0.000 983.218 0.000CO [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000H2 [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000H2O [kg/h] 0.000 0.000 0.000 160.666 30170.9 2717.9 15000.0 14601.0 14601.0 33226.5 30170.9 1910.4 30170.9CH4 [kg/h] 0.000 0.000 0.000 0.000 0.000 0.024 0.000 0.000 0.000 0.000 0.000 0.000 0.000N2 [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 3868.25 0.000H2S [kg/h] 0.000 0.000 0.000 0.000 0.000 0.006 0.000 0.000 0.000 0.000 0.000 0.000 0.000SO2 [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.322 0.000R134a [kg/h] 500.000 500.000 500.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000biomass DM [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000digestate DM [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000press fluid oDM [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 554.004 0.000 0.000 0.000press cake DM [kg/h] 0.000 0.000 0.000 1446.0 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000ash [kg/h] 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000EPH [MW] 0.002 0.003 0.007 0.001 0.125 0.003 0.051 0.132 0.050 0.113 0.064 0.177 0.140ECH

s [MW] 0.000 0.000 0.000 7.398 0.000 0.000 0.000 0.000 0.000 2.664 0.000 0.000 0.000ECH

lg [MW] 0.000 0.000 0.000 0.002 0.419 0.038 0.208 0.203 0.203 0.461 0.419 0.114 0.419

[kWel] heat losses [kWth]W1 8.0 refrigerant compressor (K17) fermenter 44.7W2 0.0 feedwater pump (K14) dewatering 70.6W3 93.3 screw press (K2) drier 75.0W4 50.8 decanter (K4) furnace 24.0W5 65.4 stirrers (fermenter + conditioning) CHP engine 65.4W6 5.2 cooler fanW7 303.1 total consumption cooler duty [kWth]W8 834.9 CHP electricity output (gross) pellet cooler 36.9W9 80.3 pellet press refrigerant condenser 26.9

252


B.3.4 Digester design and heat loss

Assumptions for the calculation of the digester dimensions are given in Table B.11. Theresulting dimensions of the digesters for ADP-3.0, ADM-3.0, ADM-3.1 and HTC-3.90 aregiven in Table B.13. Assumptions for the thermal performance are given in Table B.12.

The overall heat transfer coefficient is defined as: U =(

1α1

+ dinskins

+ 1α2

)

The heat loss of the digester is calculated as: ΔQf = U ·(

Πd2

f

4 + 2Πdf hf

2

)· (Tf − T0)

where heat loss through the ceiling is assumed to be zero, because with gas inside and airoutside, it is negligibly low.

Table B.11: Assumptions for the digester design.

active digester volume Vact = moDMOLR

, where moDM is the organic dry mattersubstrate inlet in [kg/d]

volume factor for air and fixtures Vtot/Vactive= 1.25 [221]height/diameter ratio h/d = 0.5maximum volume per unit 2000 m3

Table B.12: Assumptions for the digester heat loss calculation based on [221].

fermentation temperature Tf [°C] 37.5

outside temperature T0 [°C] 15

insulation thickness dins [m] 0.1

heat transmission coefficient, polystyrene kins [W/m/K] 0.05

heat transfer coefficient inside, agitated liquid α1 [W/m2/K] 4000

heat transfer coefficient outside, humid soil α2 [W/m2/K] 400

Table B.13: Digester design and heat loss calculation based on [221].

ADP-3.0 ADM-3.0 ADM-3.1 HTC-3.90

OLR [kgoDM/d/m3] 1.50 3.00 4.00 1.5

active digester volume Vact [m3] 8085 15104 11328 2003

total digester volume Vtot [m3] 10106 18880 14160 2503

number of digesters [-] 6 10 8 3

volume per digester Vf [m3] 1684 1888 1770 834

diameter df [m] 16.2 16.9 16.5 12.86

height hf [m] 8.1 8.4 8.3 6.43

heat loss per reactor ΔQf [kW] 7.5 8.0 7.7 4.375

heat loss total ΔQtot [kW] 44.7 80.4 61.6 13.08

253


B.3.5 Cost data

Biomethane feed-in: The cost for the transfer station and the gas pipe to the gridconnection point is shared between the biomethane producer and the grid operator [232].It is not considered in the investment cost, but a cost of 1.45 €/MWhHHV for feed-in isincluded in the annual operating cost, based on [232].

254


Tabl

eB

.14:

AD

M-3

.0-s

equi

pmen

tlis

tw

ith

inve

stm

ent

cost

s.

com

pone

ntno

.un

itX

sim

f dn

Xd/

nC

BM

/n [k€]

CB

M

[k€]

spec

ifica

tion

sco

stfu

ncti

on

dige

ster

K2

[m3]

1888

010

0%10

1888

292

2921

h=8.

4m

AD

F-1

resi

due

stor

age

tank

[m3]

9924

100%

519

8515

276

1re

side

nce

tim

e:10

0d

RST

-1

deca

nter

K13

[kg D

M/s

]0.

276

120%

10.

331

373

373

DE

C-1

pres

suri

zed

wat

er

scru

bber

K9

[mS

TP

3/h

]81

0.3

120%

197

2.3

1650

1650

PW

S-1

refr

iger

atio

n

mac

hine

K16

,K18

,

K19

[kW

th]

56.3

110%

161

.953

53co

mpr

isin

gco

mpr

esso

r,

thro

ttle

valv

e,co

nden

ser

FR

I-1

gas

cool

erK

6[m

2]

3.7

115%

14.

247

47SS

/C

SH

X-1

gas

cool

erK

8[m

2]

5.3

115%

16.

159

59SS

/C

SH

X-1

gas

cool

erK

3[m

2]

8.7

115%

110

.178

78SS

/C

SH

X-1

gas

cool

erK

10[m

2]

1.5

115%

11.

720

20C

S/

CS

HX

-1

gas

com

pres

sor

K5

[kW

]17

.711

0%1

19.4

6868

CM

P-2

mot

orK

5[k

W]

20.8

110%

122

.919

19E

M-1

gas

com

pres

sor

K7

[kW

]37

.911

0%1

41.7

107

107

CM

P-2

mot

orK

7[k

W]

43.7

110%

148

.035

35E

M-1

gas

com

pres

sor

K12

[kW

]40

.111

0%1

44.1

5454

CM

P-2

mot

orK

12[k

W]

46.1

110%

150

.737

37E

M-1

biom

ass

sizi

ng&

sort

ing

[kg/

h]66

6712

0%1

8000

162

162

shre

dder

,dru

msi

eve,

mag

neti

cse

para

tor,

air

clas

sifie

r

BSS

-1+

BSS

-2+

BSS

-3

tota

lC

BM

6445

255


Tabl

eB

.15:

AD

P-3.

0-s

equi

pmen

tlis

tw

ith

inve

stm

ent

cost

s.

com

pone

ntno

.un

itX

sim

f dn

Xd/

nC

BM

/n [k€

]

CB

M

[k€

]

spec

ifica

tion

sco

stfu

ncti

on

dige

ster

K3

[m3]

1010

610

0%6

1684

271

1626

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256

B.4 HTC plant models


B.4.1 HTC reaction

Table B.16: Mass yields and biocoal composition (d.b.) from the HTC reactor model.

sim. case HTC-3.00 HTC-3.00 HTC-3.00 HTC-3.00 HTC-3.01 HTC-3.02

feedstock wood PGW MOW EFB PGW PGW

T [°C] 220 220 220 220 210 230

t [h] 4 4 4 4 3 8

CO2 8.27% 9.14% 7.74% 9.31% 8.16% 10.55%

CO 0.38% 0.42% 0.36% 0.75% 0.38% 0.49%

CH4 0.02% 0.02% 0.02% 0.00% 0.02% 0.02%

H2 0.03% 0.04% 0.03% 0.01% 0.03% 0.04%

H2S 0.02% 0.06% 0.05% 0.00% 0.06% 0.06%

C2H4O2 3.95% 3.78% 3.20% 6.29% 3.78% 3.78%

CH2O2 0.28% 0.27% 0.23% 0.18% 0.27% 0.27%

TOMres 3.49% 3.33% 2.83% 10.71% 3.33% 3.34%

H2O 20.77% 22.97% 19.46% 13.48% 20.50% 26.51%

dissolved ash 1.37% 3.00% 3.00% 3.01% 3.00% 3.00%

biocoal 61.42% 56.98% 63.08% 56.26% 60.48% 51.94%

biocoal composition (wt %)

c 66.63% 63.00% 48.22% 65.06% 59.83% 68.31%

h 5.50% 4.69% 3.59% 6.88% 4.88% 4.37%

n 0.11% 2.00% 1.53% 0.83% 1.88% 2.19%

s 0.05% 0.18% 0.14% 0.06% 0.17% 0.20%

o 27.03% 25.56% 19.57% 24.49% 28.93% 19.93%

ash 0.68% 4.56% 26.95% 2.68% 4.30% 5.01%

Table B.17: Composition of the unspecified dissolved organic compounds from the HTCreactor model.

wood, PGW, MOW EFBc [wt%] 71.25% 66.77%h [wt%] 11.59% 11.06%n [wt%] 0.00% 1.59%s [wt%] 0.00% 0.29%o [wt%] 17.16% 20.28%HHV [MJ/kg] 36.757 34.246

257


B.4.2 Simulation data from HTC-3.00

Table B.18: HTC-3.00 flow stream data.

stream no. 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17T [°C] 15.0 90.0 56.7 184.5 58.1 219.0 232.0 152.1 54.4 214.4 203.4 186.3 184.4 152.1 204.3 204.3 56.8p [bar] 1.013 1.013 1.013 27.200 1.013 25.200 29.000 5.000 1.013 27.200 23.200 11.550 11.000 5.000 17.000 17.000 3.000m [kg/h] 6666.7 6949.8 13616.5 15878.4 196.8 15831.9 1082.8 13674.2 2237.8 16961.3 1129.4 700.0 1364.9 937.6 555.2 15276.7 13616.5H [kW] -23536 -29229 -52765 -60380 -481 -60496 -3948 -52636 -9765 -64328 -4251 -2954 -4981 -3437 -2006 -58490 -52763O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 0.00 0.00 9.97 181.18 9.98 0.00 0.00 1.59 9.97 182.77 0.00 0.60 0.01 9.37 0.61 0.00CO [kg/h] 0.00 0.00 0.00 1.93 7.90 1.95 0.00 0.00 0.54 1.93 8.42 0.00 0.42 0.03 1.49 0.46 0.00H2 [kg/h] 0.00 0.00 0.00 0.02 0.71 0.02 0.00 0.00 0.00 0.02 0.71 0.00 0.00 0.00 0.02 0.00 0.00H2O [kg/h] 4666.7 6666.7 11333.3 13571.8 6.3 14181.9 1082.8 12050.5 2225.7 14654.6 931.8 700.0 1357.0 933.3 541.2 13640.7 11333.3CH4 [kg/h] 0.00 0.00 0.00 0.20 0.27 0.21 0.00 0.00 0.09 0.20 0.34 0.00 0.08 0.02 0.12 0.10 0.00N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/h] 0.00 0.00 0.00 1.24 0.39 1.67 0.00 0.22 0.74 1.24 0.76 0.00 0.69 0.34 0.43 1.24 0.00C2H4O2 [kg/h] 0.00 95.87 95.87 105.24 0.00 176.55 0.00 164.87 8.33 105.24 4.23 0.00 5.67 3.62 2.38 174.17 95.87CH2O2 [kg/h] 0.00 6.64 6.64 7.43 0.00 12.48 0.00 11.49 0.74 7.43 0.34 0.00 0.48 0.32 0.20 12.28 6.64biocoal DM [kg/h] 0.00 0.00 0.00 0.00 0.00 1139.82 0.00 1139.82 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1139.82 0.00biomass DM [kg/h] 2000.00 0.00 2000.00 2000.00 0.00 0.00 0.00 0.00 0.00 2000.00 0.00 0.00 0.00 0.00 0.00 0.00 2000.00TOMres [kg/h] 0.00 95.08 95.08 95.08 0.00 161.78 0.00 161.78 0.00 95.08 0.00 0.00 0.00 0.00 0.00 161.78 95.08ash, diss. [kg/h] 0.00 85.53 85.53 85.53 0.00 145.53 0.00 145.53 0.00 85.53 0.00 0.00 0.00 0.00 0.00 145.53 85.53EPH [MW] 0.000 0.062 0.037 0.644 0.000 0.924 0.314 0.367 0.006 0.941 0.107 0.031 0.342 0.204 0.146 0.768 0.038ECH

s [MW] 10.062 0.000 10.062 10.062 0.000 8.158 0.000 8.158 0.000 10.062 0.000 0.000 0.000 0.000 0.000 8.158 10.062ECH

lg [MW] 0.065 1.487 1.551 1.639 0.072 2.647 0.015 2.549 0.074 1.654 0.108 0.010 0.050 0.031 0.028 2.620 1.551stream no. 20 21 22 23 27 29 30 32 34 36 37 39 40 41 42 44 46T [°C] 115.3 166.9 167.3 184.6 90.0 40.0 100.1 170.5 15.0 40.0 185.5 184.4 184.4 100.1 220.0 54.4 15.8p [bar] 9.000 9.000 15.000 15.000 1.013 1.013 1.013 1.013 1.013 1.013 27.200 11.000 11.000 1.013 25.200 1.013 33.000m [kg/h] 13958.3 15323.2 15323.2 15878.4 4083.0 1286.1 1305.2 1534.0 1227.6 2237.8 15878.4 15976.7 14611.8 388.9 15831.9 2237.8 1082.8H [kW] -53401 -58382 -58373 -60380 -17172 -1352 -4816 -1291 -33 -9801 -60361 -61054 -56072 -1435 -60476 -9765 -4784O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 88.84 284.14 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 0.60 0.60 9.97 0.00 0.00 0.01 429.15 0.00 1.59 9.97 0.61 0.01 0.00 9.98 1.59 0.00CO [kg/h] 0.01 0.43 0.43 1.93 0.00 0.00 0.02 0.00 0.00 0.54 1.93 0.46 0.04 0.01 1.95 0.54 0.00H2 [kg/h] 0.00 0.00 0.00 0.02 0.00 0.00 0.00 0.00 0.00 0.00 0.02 0.00 0.00 0.00 0.02 0.00 0.00H2O [kg/h] 11673.6 13030.6 13030.6 13571.8 3916.7 127.0 1300.2 77.2 7.8 2225.7 13571.8 14340.7 12983.7 387.4 14181.9 2225.7 1082.8CH4 [kg/h] 0.01 0.08 0.08 0.20 0.00 0.00 0.01 0.00 0.00 0.09 0.20 0.10 0.02 0.00 0.21 0.09 0.00N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 937.76 935.67 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/h] 0.12 0.81 0.81 1.24 0.00 0.00 0.38 0.00 0.00 0.74 1.24 1.24 0.56 0.11 1.67 0.74 0.00C2H4O2 [kg/h] 97.19 102.86 102.86 105.24 56.31 1.83 4.11 0.00 0.00 8.33 105.24 174.17 168.50 1.22 176.55 8.33 0.00CH2O2 [kg/h] 6.75 7.23 7.23 7.43 3.90 0.13 0.40 0.00 0.00 0.74 7.43 12.28 11.80 0.12 12.48 0.74 0.00biocoal DM [kg/h] 0.00 0.00 0.00 0.00 0.00 1139.82 0.00 0.00 0.00 0.00 0.00 1139.82 1139.82 0.00 1139.82 0.00 0.00biomass DM [kg/h] 2000.00 2000.00 2000.00 2000.00 0.00 0.00 0.00 0.00 0.00 0.00 2000.00 0.00 0.00 0.00 0.00 0.00 0.00TOMres [kg/h] 95.08 95.08 95.08 95.08 55.86 7.59 0.00 0.00 0.00 0.00 95.08 161.78 161.78 0.00 161.78 0.00 0.00ash, diss. [kg/h] 85.53 85.53 85.53 85.53 50.25 9.75 0.00 0.00 0.00 0.00 85.53 145.53 145.53 0.00 145.53 0.00 0.00EPH [MW] 0.207 0.498 0.503 0.640 0.037 0.001 0.201 0.016 0.000 0.003 0.651 0.933 0.591 0.060 0.933 0.006 0.001ECH

s [MW] 10.062 10.062 10.062 10.062 0.000 8.158 0.000 0.000 0.000 0.000 10.062 8.158 8.158 0.000 8.158 0.000 0.000ECH

lg [MW] 1.563 1.612 1.612 1.639 0.874 0.088 0.038 0.043 0.002 0.074 1.639 2.630 2.580 0.011 2.647 0.074 0.015stream no. 54 55 56 57 58 59 60 61 62 63 64 67 68 70 71 72 73T [°C] 90.0 160.5 15.0 151.4 152.1 15.0 42.7 100.1 15.0 43.6 94.8 60.0 43.6 15.0 75.3 800.0 100.1p [bar] 1.013 32.000 1.000 5.000 5.000 1.013 1.013 1.013 1.013 1.013 1.013 3.000 1.013 1.013 3.000 1.013 1.013m [kg/h] 11032.7 1082.8 1082.8 595.8 14269.9 1168.8 1286.1 748.8 25500.0 748.8 27034.0 196.8 27676.7 117.3 13616.5 7.6 37.9H [kW] -46401 -4604 -4785 -2534 -55169 -1243 -1351 -2763 -677 -3279 -1425 -481 -4194 -125 -52505 0 -140O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 5902.36 0.00 5991.20 0.00 5991.20 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 0.00 0.00 0.01 0.01 0.00 0.00 0.00 0.00 0.00 429.15 181.18 429.15 0.00 0.00 0.00 0.00CO [kg/h] 0.00 0.00 0.00 0.02 0.02 0.00 0.00 0.01 0.00 0.01 0.00 7.90 0.00 0.00 0.00 0.00 0.00H2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.71 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 10583.3 1082.8 1082.8 593.0 12643.5 115.4 127.0 746.0 161.4 746.0 238.6 6.3 871.5 11.6 11333.3 0.00 37.8CH4 [kg/h] 0.00 0.00 0.00 0.01 0.01 0.00 0.00 0.01 0.00 0.01 0.00 0.27 0.00 0.00 0.00 0.00 0.00N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 19436.2 0.00 20374.0 0.00 20374.0 0.00 0.00 0.00 0.00H2S [kg/h] 0.00 0.00 0.00 0.22 0.43 0.00 0.00 0.22 0.00 0.22 0.00 0.39 0.06 0.00 0.00 0.00 0.01C2H4O2 [kg/h] 152.14 0.00 0.00 2.30 167.17 0.00 1.83 2.36 0.00 2.36 0.00 0.00 9.10 0.00 95.87 0.00 0.12CH2O2 [kg/h] 10.54 0.00 0.00 0.20 11.69 0.00 0.13 0.23 0.00 0.23 0.00 0.00 0.63 0.00 6.64 0.00 0.01biocoal DM [kg/h] 0.00 0.00 0.00 0.00 1139.82 1053.40 1139.82 0.00 0.00 0.00 0.00 0.00 0.00 105.71 0.00 0.00 0.00biomass DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 2000.00 0.00 0.00TOMres [kg/h] 150.94 0.00 0.00 0.00 161.78 0.00 7.59 0.00 0.00 0.00 0.00 0.00 0.00 0.00 95.08 0.00 0.00ash, diss. [kg/h] 135.78 0.00 0.00 0.00 145.53 0.00 9.75 0.00 0.00 0.00 0.00 0.00 0.00 0.00 85.53 5.63 0.00EPH [MW] 0.099 0.036 0.000 0.017 0.384 0.000 0.001 0.116 0.000 0.001 0.072 0.004 0.024 0.000 0.076 0.002 0.006ECH

s [MW] 0.000 0.000 0.000 0.000 8.158 7.488 8.158 0.000 0.000 0.000 0.000 0.000 0.000 0.751 10.062 0.000 0.000ECH

lg [MW] 2.361 0.015 0.015 0.020 2.569 0.002 0.088 0.022 0.032 0.022 0.051 0.072 0.098 0.000 1.551 0.019 0.001stream no. 76 78 80 81 82 85 86 87 90 91 92 93 94 95 96 99 100T [°C] 100.0 43.6 218.3 15.0 15.3 90.1 170.5 114.2 90.1 90.0 100.1 41.6 40.0 60.0 100.1 42.7 15.0p [bar] 3.000 1.013 23.200 1.013 11.850 1.050 23.200 23.200 1.013 32.700 1.013 1.013 1.013 3.000 1.013 1.013 1.013m [kg/h] 13616.5 37.92 1129.35 700.00 700.00 25500.0 1129.35 1129.35 1227.59 1082.81 129.60 129.60 4082.98 932.59 12964.8 3.25 1286.09H [kW] -52155 -166 -3861 -3094 -3093 -133 -4390 -4480 -6 -4694 -478 -568 -17396 -4059 -50354 -3 -1368O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 5902.36 0.00 0.00 284.14 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 0.00 182.77 0.00 0.00 0.00 182.77 182.77 0.00 0.00 0.00 0.00 0.00 1.59 0.00 0.00 0.00CO [kg/h] 0.00 0.00 8.42 0.00 0.00 0.00 8.42 8.42 0.00 0.00 0.00 0.00 0.00 0.52 0.00 0.00 0.00H2 [kg/h] 0.00 0.00 0.71 0.00 0.00 0.00 0.71 0.71 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 11333.3 37.78 931.78 700.00 700.00 161.44 931.78 931.78 7.77 1082.81 129.11 129.11 3916.66 925.48 11343.2 0.00 126.98CH4 [kg/h] 0.00 0.00 0.34 0.00 0.00 0.00 0.34 0.34 0.00 0.00 0.00 0.00 0.00 0.08 0.00 0.00 0.00N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 19436.2 0.00 0.00 935.67 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/h] 0.00 0.01 0.76 0.00 0.00 0.00 0.76 0.76 0.00 0.00 0.04 0.04 0.00 0.36 0.06 0.00 0.00C2H4O2 [kg/h] 95.87 0.12 4.23 0.00 0.00 0.00 4.23 4.23 0.00 0.00 0.41 0.41 56.31 4.22 163.07 0.00 0.00CH2O2 [kg/h] 6.64 0.01 0.34 0.00 0.00 0.00 0.34 0.34 0.00 0.00 0.04 0.04 3.90 0.34 11.29 0.00 0.00biocoal DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1139.82 0.00 1159.11biomass DM [kg/h] 2000.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00TOMres [kg/h] 95.08 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 55.86 0.00 161.78 3.25 0.00ash, diss. [kg/h] 85.53 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 50.25 0.00 145.53 0.00 0.00EPH [MW] 0.146 0.000 0.266 0.000 0.000 0.082 0.054 0.026 0.003 0.011 0.020 0.000 0.005 0.003 0.140 0.000 0.000ECH

s [MW] 10.062 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 8.158 0.000 8.240ECH

lg [MW] 1.551 0.001 0.108 0.010 0.010 0.032 0.108 0.108 0.002 0.015 0.004 0.004 0.874 0.036 2.531 0.034 0.002

[kWel] [kWel] heat losses [kWth] cooler duty [kWth]W1 2.9 slurry pump (K2) W8 31.8 drier fan (K38) biomass slurry (K11) 18.8 reactor offgas (K42) 60.1W2 9.3 slurry pump (K6) W9 64.3 pellet press (K46) biocoal slurry (K16) 20.5 condenser (K28) 0.0W3 2.0 pump (K27) W10 412.8 total plant boiler (K25) 22.9 condensate (K24) 35.8W4 10.7 slurry pump (K8) W11 0.5 pump (K40) HTC reactor (K14) 8.1 waste water (K33) 224.2W5 22.9 slurry pump (K10) W13 12.5 cooler fans drier (K22) 32.1 filter press (K21) 138.9W7 13.1 filter press (K21) W14 242.9 waste water aeration

258


B.4.3 Simulation data from HTC-5.00

biocoal

condensate

reactor

pelletpress

flue gas

filter pressair

76

50

14

47

19 20 21 22 23

6

15

13

96

54

53

1639

37 4

10

7

67

8011 86

12

87

82

81

56

62

92

41

97

61

28

63

78

64

70

60

29 75

90

7873

63

95

95

93

93

9

32

34

46

91

55

8340

8

57

44

36

5868+99

30

42

K2 K3 K4 K5

K28

K24

K45

K6 K7 K8 K9 K10 K11

K16 K15

K29

K19

K31

K21 K37

K46K22

K20

K18

K41

K40

K36

K27

K12

K32

K26

K25

K42

K43

K13

K34

K17

17 71

air

drier

A

A

BC

C

D

D

B

EFBfluegas

27

1 K1

K33

3

2W1 W2 W4

W9W7

W5

W11

W8

W3

shells,fibres

wastewater

aeration

discharge

94

W14

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

K44

ash72

Figure B.1: Flowsheet of HTC-5.00.

Table B.19: Composition (d.b.) and heating value of fibres and shells used as boiler fuel.

c 49.18%h 7.00%n 0.64%s 0.07%o 38.59%a 4.52%w (w.b.) 32.50%HHVdry [MJ/kg] 21.332

It is assumed that the boiler fuel comprises 75% fibres with a water content of 40% plus25% shells with a water content of 10%. The dry matter composition is assumed to beidentical to that of the EFB.

259



stream no. 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 19T [°C] 27.0 90.0 67.2 176.9 58.1 219.0 240.9 152.2 72.9 206.2 197.1 186.3 184.5 152.2 203.4 203.4 67.3 124.3p [bar] 1.0 1.0 1.0 32.2 1.0 30.2 34.0 5.0 1.0 32.2 28.2 11.6 11.0 5.0 17.0 17.0 3.0 3.0m [kg/h] 5714.3 8449.0 14163.3 16234.3 206.8 16755.1 1019.3 14493.4 1620.0 17253.5 498.4 200.0 855.9 969.8 636.0 16119.1 14163.3 14742.4H [kW] -18983 -34014 -52997 -59981 -499 -62176 -3714 -53995 -7026 -63695 -1660 -844 -3112 -3548 -2254 -59922 -52995 -54641O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 0.00 0.00 32.33 185.61 32.35 0.00 0.00 0.52 32.33 186.11 0.00 1.75 0.04 30.55 1.80 0.00 0.03CO [kg/h] 0.00 0.00 0.00 10.93 14.57 11.06 0.00 0.02 0.42 10.93 14.87 0.00 2.12 0.28 8.65 2.42 0.00 0.17H2 [kg/h] 0.00 0.00 0.00 0.02 0.19 0.02 0.00 0.00 0.00 0.02 0.19 0.00 0.00 0.00 0.02 0.00 0.00 0.00H2O [kg/h] 3714.29 7619.05 11333.33 13342.11 6.36 14336.76 1019.28 12140.82 1606.40 14361.39 294.21 200.00 844.02 961.01 590.92 13745.84 11333.33 11907.18CH4 [kg/h] 0.00 0.00 0.00 0.09 0.04 0.10 0.00 0.00 0.01 0.09 0.05 0.00 0.03 0.01 0.05 0.04 0.00 0.01N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00C2H4O2 [kg/h] 0.00 241.13 241.13 259.32 0.01 382.26 0.00 360.78 12.22 259.32 2.92 0.00 7.70 8.18 5.60 376.66 241.13 246.02CH2O2 [kg/h] 0.00 6.68 6.68 7.29 0.00 10.79 0.00 10.06 0.45 7.29 0.09 0.00 0.26 0.29 0.18 10.61 6.68 6.85biocoal DM [kg/h] 0.00 0.00 0.00 0.00 0.00 1125.26 0.00 1125.26 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1125.26 0.00 0.00biomass DM [kg/h] 2000.00 0.00 2000.00 2000.00 0.00 0.00 0.00 0.00 0.00 2000.00 0.00 0.00 0.00 0.00 0.00 0.00 2000.00 2000.00TOMres [kg/h] 0.00 454.35 454.35 454.35 0.00 668.46 0.00 668.46 0.00 454.35 0.00 0.00 0.00 0.00 0.00 668.46 454.35 454.35ash, diss. [kg/h] 0.00 127.80 127.80 127.80 0.00 188.03 0.00 188.03 0.00 127.80 0.00 0.00 0.00 0.00 0.00 188.03 127.80 127.80stream no. 20 21 22 23 27 29 30 32 34 36 37 39 40 41 42 44 46 53T [°C] 124.7 156.4 156.8 176.3 90.0 40.0 100.2 130.0 27.0 40.0 177.9 184.5 184.5 100.2 220.0 72.9 27.9 90.0p [bar] 9.0 9.0 15.0 15.0 1.0 1.0 1.0 1.0 1.0 1.0 32.2 11.0 11.0 1.0 30.2 1.0 38.0 1.0m [kg/h] 14742.4 15598.3 15598.3 16234.3 3149.6 1300.6 1328.4 1728.9 1331.9 1620.0 16234.3 16319.1 15463.2 372.5 16755.1 1620.0 1019.3 1957.1H [kW] -54634 -57746 -57737 -59991 -12680 -1481 -4894 -1717 -88 -7086 -59963 -60655 -57543 -1372 -62154 -7026 -4490 -4209O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 101.82 304.72 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.03 1.78 1.78 32.33 0.00 0.00 0.02 442.51 0.00 0.52 32.33 1.80 0.05 0.01 32.35 0.52 0.00 0.00CO [kg/h] 0.17 2.29 2.29 10.93 0.00 0.00 0.13 0.00 0.00 0.42 10.93 2.42 0.29 0.04 11.06 0.42 0.00 0.00H2 [kg/h] 0.00 0.00 0.00 0.02 0.00 0.00 0.00 0.00 0.00 0.00 0.02 0.00 0.00 0.00 0.02 0.00 0.00 0.00H2O [kg/h] 11907.18 12751.19 12751.19 13342.11 2840.21 127.26 1318.56 180.13 23.79 1606.40 13342.11 13945.84 13101.83 369.73 14336.76 1606.40 1019.28 750.17CH4 [kg/h] 0.01 0.04 0.04 0.09 0.00 0.00 0.01 0.00 0.00 0.01 0.09 0.04 0.01 0.00 0.10 0.01 0.00 0.00N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1004.28 1003.42 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00C2H4O2 [kg/h] 246.02 253.72 253.72 259.32 89.89 4.03 9.31 0.00 0.00 12.22 259.32 376.66 368.96 2.61 382.26 12.22 0.00 23.74CH2O2 [kg/h] 6.85 7.11 7.11 7.29 2.49 0.11 0.36 0.00 0.00 0.45 7.29 10.61 10.35 0.10 10.79 0.45 0.00 0.66biocoal DM [kg/h] 0.00 0.00 0.00 0.00 0.00 1125.26 0.00 0.00 0.00 0.00 0.00 1125.26 1125.26 0.00 1125.26 0.00 0.00 1125.26biomass DM [kg/h] 2000.00 2000.00 2000.00 2000.00 0.00 0.00 0.00 0.00 0.00 0.00 2000.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00TOMres [kg/h] 454.35 454.35 454.35 454.35 169.37 31.32 0.00 0.00 0.00 0.00 454.35 668.46 668.46 0.00 668.46 0.00 0.00 44.74ash, diss. [kg/h] 127.80 127.80 127.80 127.80 47.64 12.58 0.00 0.00 0.00 0.00 127.80 188.03 188.03 0.00 188.03 0.00 0.00 12.58stream no. 54 55 56 57 58 60 61 62 63 64 67 68 70 71 72 73 75 76T [°C] 90.0 120.0 27.0 149.5 152.1 44.6 100.2 27.0 66.0 90.2 60.0 46.1 27.0 84.3 800.0 100.2 27.0 100.1p [bar] 1.0 37.0 1.0 5.0 5.0 1.0 1.0 1.0 1.0 1.1 3.0 1.0 1.0 3.0 1.0 1.0 1.0 3.0m [kg/h] 11598.6 1019.3 1019.3 390.7 14884.2 1300.6 798.0 31100.0 798.0 31100.0 206.8 31743.2 198.2 14163.3 8.0 36.0 1300.6 14163.3H [kW] -46694 -4383 -4491 -1659 -55654 -1479 -2940 -2058 -3469 -1494 -499 -4248 -459 -52753 0 -132 -1489 -52523O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 7115.03 0.00 7115.03 0.00 7115.03 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 0.00 0.00 0.02 0.02 0.00 0.01 0.00 0.01 0.00 185.61 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO [kg/h] 0.00 0.00 0.00 0.11 0.13 0.00 0.08 0.00 0.08 0.00 14.57 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.19 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 10459.25 1019.28 1019.28 387.16 12527.98 127.26 792.05 555.49 792.05 555.49 6.36 1178.41 64.42 11333.33 0.00 35.68 127.26 11333.33CH4 [kg/h] 0.00 0.00 0.00 0.00 0.01 0.00 0.00 0.00 0.00 0.00 0.04 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 23429.48 0.00 23429.48 0.00 23429.48 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00C2H4O2 [kg/h] 331.02 0.00 0.00 3.29 364.08 4.03 5.59 0.00 5.59 0.00 0.01 19.71 0.00 241.13 0.00 0.25 4.03 241.13CH2O2 [kg/h] 9.16 0.00 0.00 0.12 10.18 0.11 0.22 0.00 0.22 0.00 0.00 0.55 0.00 6.68 0.00 0.01 0.11 6.68biocoal DM [kg/h] 0.00 0.00 0.00 0.00 1125.26 1125.26 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1125.26 0.00biomass DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 133.80 2000.00 0.00 0.00 0.00 2000.00TOMres [kg/h] 623.73 0.00 0.00 0.00 668.46 31.32 0.00 0.00 0.00 0.00 0.00 0.00 0.00 454.35 0.00 0.00 31.32 454.35ash, diss. [kg/h] 175.45 0.00 0.00 0.00 188.03 12.58 0.00 0.00 0.00 0.00 0.00 0.00 0.00 127.80 6.05 0.00 12.58 127.80stream no. 78 80 81 82 86 87 90 91 92 93 94 95 96 99T [°C] 58.7 218.7 27.0 27.3 177.1 137.6 90.2 90.0 100.2 100.0 40.0 60.0 100.2 44.6p [bar] 1.0 28.2 1.0 11.9 28.2 28.2 1.0 37.7 1.0 1.0 1.0 3.0 1.0 1.0m [kg/h] 36.0 498.4 200.0 200.0 498.4 498.4 1331.9 1019.3 122.0 122.0 3149.6 291.6 13555.8 13.4H [kW] -157 -1548 -881 -881 -1697 -1732 -64 -4419 -449 -520 -12847 -1266 -50761 -12O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 304.72 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 186.11 0.00 0.00 186.11 186.11 0.00 0.00 0.00 0.00 0.00 0.50 0.00 0.00CO [kg/h] 0.00 14.87 0.00 0.00 14.87 14.87 0.00 0.00 0.01 0.01 0.00 0.29 0.00 0.00H2 [kg/h] 0.00 0.19 0.00 0.00 0.19 0.19 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 35.68 294.21 200.00 200.00 294.21 294.21 23.79 1019.28 121.09 121.09 2840.21 287.85 11209.43 0.00CH4 [kg/h] 0.00 0.05 0.00 0.00 0.05 0.05 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 1003.42 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00C2H4O2 [kg/h] 0.25 2.92 0.00 0.00 2.92 2.92 0.00 0.00 0.86 0.86 89.89 2.91 354.76 0.00CH2O2 [kg/h] 0.01 0.09 0.00 0.00 0.09 0.09 0.00 0.00 0.03 0.03 2.49 0.09 9.82 0.00biocoal DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1125.26 0.00biomass DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00TOMres [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 169.37 0.00 668.46 13.42ash, diss. [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 47.64 0.00 188.03 0.00

[kWel] [kWel] heat losses [kWth] cooler duty [kWth]W1 3.2 slurry pump (K2) W8 40.4 drier fan (K38) biomass slurry (K11) 18.6 reactor offgas (K42) 33.4W2 10.5 slurry pump (K6) W9 65.0 pellet press (K46) biocoal slurry (K16) 21.4 condenser (K28) 0.0W3 2.2 pump (K27) W10 765.0 total plant boiler (K25) 25.3 condensate (K24) 59.4W4 11.4 slurry pump (K8) W11 0.1 pump (K40) HTC reactor (K14) 7.5 waste water (K33) 166.9W5 36.5 slurry pump (K10) W13 4.2 cooler fans drier (K22) 35.1 filter press (K21) 142.5W7 13.4 filter press (K21) W14 578.0 waste water aeration

260


B.4.4 The fate of nutrients from EFB in HTC-5.00

Table B.21 shows the nutrients in EFB and the fraction that ends up in the HTC wastewater according to lab-scale experiments [23], as well as the GHG emissions that wouldbe caused by the production and application of mineral fertilizers with the same nutrientvalue. The specific GHG emissions per kg fertilizer are based on [405].

Table B.22 shows the fertilizer value of the EFB calculated with fertilizer prices for Malay-sia and Indonesia.

Table B.21: Nutrients in EFB, HTC waste water, and GHG emissions of the equivalentmineral fertilizers, per t EFB (DM).

kg element % inwastewater

kg fertilizerequivalent

kg CO2eq/ kgfertilizer

kg CO2eq

totalkg CO2eq

in wastewater

N 6.367 23% 1.671 N1) 3.29 5.499 1.26523% 4.696 N 2) 2.68 12.584 2.894

P 0.513 73% 1.114 P2O5 2.46 2.742 2.001K 7.257 88% 12.095 K2O 0.5 6.047 5.322Ca 1.727 72%Mg 1.282 88%

total 26.872 11.4821) in urea2) in ammonia sulphate

Table B.22: Monetary value of the nutrients in the EFB and in the HTC waste water,calculated with fertilizer prices (1) for Malaysia [64] and (2) Indonesia [406], per t EFB(DM).

fertilizer price 1 value, total value, in liquid fertilizer price 2 value, total value, in liquid[€/kg] [€ ] [€ ] [€/kg] [€ ] [€ ]

N 0.377 2.40 0.551 0.524 3.339 0.768P 0.366 0.19 0.137 1.078 0.553 0.403K 0.124 0.90 0.790 0.819 5.944 5.231Ca 0.000 0.345 0.596 0.429Mg 0.715 0.92 0.807 0.969 1.242 1.093

total 4.40 2.29 11.67 7.92

261


B.4.5 Cost data from the HTC base design

Table B.23: Investment cost from the HTC base cases.

1.00-s 1.00-m 2.00-s 2.00-m 3.00-s 3.00-m 4.00-s 4.00-m 5.00-s 5.00-m1)

[M€] [M€] [M€] [M€] [M€] [M€] [M€] [M€] [M€] [M€]

HTC reactor 1.90 8.34 2.03 9.07 1.77 7.59 1.78 7.63 1.98 8.87

slurry pumps 0.51 1.49 0.55 1.59 0.47 1.37 0.46 1.35 0.57 1.67

flash tanks 0.54 1.74 0.54 1.73 0.53 1.70 0.53 1.72 0.54 1.73

filter press 0.75 2.33 0.75 2.33 0.74 2.30 0.75 2.33 0.74 2.32

biocoal drier (incl. HX) 0.81 2.01 0.78 1.92 0.78 1.93 0.82 2.04 0.78 1.93

pellet press 0.36 1.06 0.34 1.02 0.34 1.00 0.36 1.07 0.34 1.01

boiler 0.23 0.82 0.23 0.81 0.23 0.84 0.28 1.00 0.25 0.91

heat exchangers 0.47 0.86 0.48 0.89 0.51 0.91 0.50 0.90 0.43 0.80

coolers 0.31 0.68 0.33 0.70 0.35 0.73 0.34 0.72 0.35 0.73

pumps and fans 0.04 0.07 0.04 0.07 0.04 0.07 0.04 0.07 0.04 0.07

biomass slurry preparation 0.08 0.24 0.09 0.25 0.09 0.25 0.09 0.25 0.09 0.25

biocoal storage & handling 0.18 0.51 0.17 0.49 0.17 0.48 0.18 0.51 0.17 0.48

biomass sizing & sorting 0.00 0.00 0.13 0.32 0.16 0.44 0.16 0.44 0.11 0.33

waste water treatment 0.23 0.59 0.23 0.60 0.34 0.88 0.34 0.87 0.29 0.75

total CBM 6.41 20.74 6.67 21.79 6.52 20.50 6.64 20.92 6.68 21.85

offsite cost 1.18 3.40 1.08 3.12 1.08 3.12 0.95 2.75 0.67 1.90

fees & contingencies 0.96 3.11 1.00 3.27 0.98 3.08 1.00 3.14 1.00 3.28

start-up 0.31 0.97 0.30 0.88 0.30 0.86 0.29 0.85 0.23 0.82

working capital 0.68 2.45 0.43 1.01 0.43 1.04 0.41 0.98 0.19 0.80

AFUDC 0.85 2.73 0.87 2.82 0.86 2.67 0.86 2.68 0.83 2.70

residual value (NPV) -0.12 -0.44 -0.05 -0.07 -0.05 -0.10 -0.04 -0.06 0.05 0.13

TCI 10.27 32.97 10.30 32.81 10.10 31.17 10.10 31.27 9.66 31.481) The investment costs for the plant equipment were estimated for HTC-5.00-m with a processing capacity of 59.3

MWHHV of EFB. The TCI for HTC-5.00-m* with a capacity of 26.6 MWHHV of EFB were then estimated with

Equation 3.11. See subsection 4.6.5.

262


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264


B.4.6 Exergy analysis and exergoeconomic analysis

Table B.25: Component results from the exergoeconomic analysis of HTC-3.00-s.

No type ED EL yD yD* EF EP cF cP r f CD Z CD+Z[kW] [kW] [%] [%] [kW] [kW] [%] [€/GJ] [€/GJ] [-] [-] [€/h] [€/h] [€/h]

1 slurry-mix 25.8 0.0 0.5% 1.3% 44.1 18.3 41.4% 52.1 253.7 387% 63% 4.84 8.41 13.252 slurry pump 2.1 0.0 0.0% 0.1% 2.9 0.8 29.0% 28.7 655.4 2186% 89% 0.21 1.68 1.903 slurry preheater HX 20.2 0.0 0.4% 1.0% 58.2 38.0 65.3% 142.0 250.5 76% 30% 10.34 4.51 14.854 slurry preheater HX 43.2 0.0 0.8% 2.2% 112.8 69.5 61.7% 128.3 222.7 74% 16% 19.97 3.68 23.655 steam / slurry mix 16.9 0.0 0.3% 0.9% 69.6 52.6 75.7% 121.7 160.8 32% 0% 7.41 0.00 7.416 slurry pump 5.8 0.0 0.1% 0.3% 9.3 3.5 37.3% 28.7 423.5 1377% 88% 0.60 4.31 4.917 steam / slurry mix 50.0 0.0 0.9% 2.5% 297.3 247.2 83.2% 110.1 132.4 20% 0% 19.82 0.00 19.828 slurry pump 6.0 0.0 0.1% 0.3% 10.7 4.7 44.1% 28.7 365.6 1175% 89% 0.61 5.08 5.709 steam / slurry mix 10.1 0.0 0.2% 0.5% 124.9 114.8 91.9% 144.6 157.3 9% 0% 5.28 0.00 5.2810 slurry pump 12.3 0.0 0.2% 0.6% 22.9 10.6 46.3% 28.7 332.8 1061% 89% 1.27 10.34 11.6011 heat loss 7.2 0.0 0.1% 0.4%12 air preheater 2.9 0.0 0.1% 0.1% 5.8 2.9 50.3% 142.0 438.8 209% 53% 1.47 1.64 3.1113 steam / slurry mix 16.7 0.0 0.3% 0.8% 254.2 237.5 93.4% 14.2 15.2 7% 0% 0.85 0.00 0.8514 HTC reactor 546.2 0.0 10.4% 27.6% 3226.7 2680.6 83.1% -2.7 3.0 -210% 91% 5.85 60.13 65.9815 evaporator (offgas) 14.2 0.0 0.3% 0.7% 158.3 144.2 91.1% 33.6 42.5 26% 63% 1.71 2.91 4.6216 heat loss 8.8 0.0 0.2% 0.4%17 flash 1 8.4 0.0 0.2% 0.4% 121.9 113.5 93.1% 128.2 155.4 21% 65% 3.88 7.24 11.1218 flash 2 1) 0.0 0.0 0.0% 0.0% 261.9 261.9 100.0% 113.4 118.8 5% 100% 0.00 5.07 5.0719 flash 3 19.4 0.0 0.4% 1.0% 185.8 166.4 89.6% 113.4 131.7 16% 28% 7.91 3.06 10.9720 flash 4 43.0 0.0 0.8% 2.2% 209.2 166.2 79.5% 114.1 147.9 30% 13% 17.64 2.58 20.2321 mech. dewatering 43.0 0.0 0.8% 2.2% 28.7 89% 4.44 35.54 39.9822 drier 59.0 0.0 1.1% 3.0% 238.4 31% 50.60 22.51 73.1124 condensate cooler 3.5 0.0 0.1% 0.2% 90.3 70% 1.14 2.65 3.7925 boiler 470.8 0.0 8.9% 23.8% 750.6 278.3 37.1% 5.9 23.7 303% 44% 9.95 7.88 17.8326 FWPH 3.1 0.0 0.1% 0.2% 28.1 24.9 88.8% 33.6 64.8 92% 86% 0.38 2.41 2.7927 pump 1.0 0.0 0.0% 0.1% 2.0 1.0 47.9% 28.7 354.8 1138% 90% 0.11 1.03 1.1328 condenser 0.0 0.0 0.0% 0.0%29 mix 10.9 0.0 0.2% 0.6% 943.7 932.8 98.8% 112.0 113.3 1% 0% 4.40 0.00 4.4031 mix 0.0 0.0 0.0% 0.0% 384.4 384.4 100.0% 114.0 114.0 0% 0% 0.02 0.00 0.0232 drier HX 54.1 0.0 1.0% 2.7% 114.4 60.4 52.7% 142.0 286.3 102% 13% 27.64 3.95 31.5933 waste water cooler 32.2 0.0 0.6% 1.6% 32.2 114.0 21% 13.21 3.60 16.8136 FWPH 9.9 0.0 0.2% 0.5% 19.8 10.0 50.3% 142.0 307.8 117% 15% 5.03 0.92 5.9637 biocoal cooler 0.1 0.0 0.0% 0.0% 0.1 114.0 85% 0.05 0.30 0.3638 drier fan 10.5 0.0 0.2% 0.5% 31.8 21.3 67.1% 28.7 44.2 54% 10% 1.08 0.12 1.2039 drier gas mix 49.8 0.0 0.9% 2.5% 172.8 123.0 71.2% 109.0 153.1 40% 0% 19.54 0.00 19.5440 pump 0.3 0.0 0.0% 0.0% 0.5 0.2 45.7% 28.7 375.7 1210% 90% 0.03 0.24 0.2641 FWPH (offgas) 22.4 0.0 0.4% 1.1% 53.2 30.8 57.9% 33.6 70.6 110% 34% 2.71 1.39 4.1042 offgas cooler 19.0 0.0 0.4% 1.0% 91.8 72.2 78.6% 4.8 11.1 133% 81% 0.33 1.40 1.7243 boiler fuel valve 3.7 0.0 0.1% 0.2% 76.2 72.5 95.2% 12.3 12.9 5% 0% 0.16 0.00 0.1644 heat loss storage 0.6 0.0 0.0% 0.0% 114.0 0% 0.24 0.00 0.2445 condensate collector 0.4 0.0 0.0% 0.0% 80.7 80.3 99.5% 4.9 4.9 -1% 0% 0.01 0.00 0.0146 pellet press 68.3 0.0 1.3% 3.5% 28.7 69% 7.05 15.35 22.3947 misc. el. consumption 255.4 0.0 4.8% 12.9% 28.7 0% 26.36 0.00 26.36

total plant 1976.9 1108.2 37.5% 100% 5273.5 2188.6 41.5%

1) exergy destruction takes place in upstream mixer K29

265


Table B.26: Flow stream results from the exergoeconomic analysis of HTC-3.00-s.

ECHS cCH

S CCHS ECH

lg cCHlg CCH

lg EPH cPH CPH ETOT cTOT CTOT W c C[kW] [€/GJ] [€/h] [kW] [€/GJ] [€/h] [kW] [€/GJ] [€/h] [kW] [€/GJ] [€/h] [kW] [€/GJ] [€/h]

1 10062 -2.71 -98.29 64.8 0.00 0.00 0.0 0.00 0.00 10127.1 -2.70 -98.29 W1 2.9 28.67 0.302 0 0.00 0.00 1487.2 -2.91 -15.59 62.5 114.01 25.65 1549.7 1.80 10.06 W2 9.3 28.67 0.963 10062 -2.71 -98.29 1551.3 -2.79 -15.58 37.3 253.70 34.06 11651.0 -1.90 -79.81 W3 2.0 28.67 0.214 10062 -2.71 -98.29 1638.7 -3.08 -18.18 643.7 176.53 409.07 12344.8 6.58 292.60 W4 10.7 28.67 1.105 0 0.00 0.00 72.2 12.86 3.34 0.3 33.65 0.04 72.5 12.95 3.38 W5 22.9 28.67 2.366 8158 -1.02 -30.07 2646.6 -2.92 -27.84 924.1 126.68 421.46 11728.4 8.61 363.55 W6 45.7 28.67 4.727 0 0.00 0.00 15.0 0.00 0.00 314.2 36.98 41.82 329.2 35.29 41.82 W7 13.1 28.67 1.358 8158 -1.02 -30.07 2549.2 -2.91 -26.72 367.3 113.35 149.89 11074.2 2.34 93.10 W8 31.8 28.67 3.289 0 0.00 0.00 74.2 -2.12 -0.57 6.1 90.32 1.97 80.3 4.87 1.41 W9 64.3 28.67 6.64

10 10062 -2.71 -98.29 1653.7 -3.05 -18.18 941.3 133.06 450.90 12657.3 7.34 334.43 W10 412.8 28.67 42.6011 0 0.00 0.00 108.0 -3.05 -1.19 107.5 33.65 13.02 215.5 15.25 11.83 W11 0.5 28.67 0.0512 0 0.00 0.00 9.7 0.00 0.00 31.0 72.66 8.12 40.7 55.33 8.12 W12 0.0 0.00 0.0013 0 0.00 0.00 49.8 -2.91 -0.52 341.6 117.52 144.50 391.4 102.18 143.98 W13 12.5 28.67 1.2914 0 0.00 0.00 31.1 -2.91 -0.33 204.4 128.29 94.39 235.4 110.97 94.06 W14 242.9 28.67 25.0615 0 0.00 0.00 27.8 -2.92 -0.29 145.9 149.05 78.27 173.6 124.75 77.9816 8158 -1.02 -30.07 2620.3 -2.92 -27.57 768.4 126.68 350.44 11546.4 7.04 292.8117 10062 -2.71 -98.29 1551.3 -2.79 -15.58 38.1 262.55 36.04 11651.8 -1.86 -77.8319 10062 -2.71 -98.29 1562.7 -2.79 -15.70 203.3 219.29 160.49 11828.3 1.09 46.5020 10062 -2.71 -98.29 1562.7 -2.79 -15.70 206.7 222.71 165.76 11831.7 1.22 51.7721 10062 -2.71 -98.29 1612.4 -2.98 -17.29 498.4 173.52 311.32 12173.1 4.47 195.7522 10062 -2.71 -98.29 1612.4 -2.98 -17.29 503.1 175.31 317.51 12177.8 4.61 201.9323 10062 -2.71 -98.29 1638.7 -3.08 -18.18 640.2 171.97 396.38 12341.3 6.30 279.9127 0 0.00 0.00 873.6 -2.91 -9.16 36.7 114.01 15.07 910.4 1.80 5.9129 8158 3.14 92.20 88.0 -2.91 -0.92 0.6 114.01 0.24 8246.3 3.08 91.5230 0 0.00 0.00 38.5 -2.91 -0.40 201.4 141.96 102.93 239.8 118.74 102.5232 0 0.00 0.00 42.9 12.86 1.99 16.3 3.51 0.21 59.2 10.28 2.1934 0 0.00 0.00 1.5 0.00 0.00 0.0 0.00 0.00 1.5 0.00 0.0036 0 0.00 0.00 74.2 0.00 0.00 2.6 90.32 0.84 76.8 3.02 0.8437 10062 -2.71 -98.29 1638.7 -3.08 -18.18 650.8 174.59 409.07 12351.9 6.58 292.6039 8158 -1.02 -30.07 2629.9 -2.91 -27.56 932.8 113.35 380.64 11720.4 7.66 323.0140 8158 -1.02 -30.07 2580.2 -2.91 -27.04 591.1 113.35 241.22 11329.0 4.51 184.1141 0 0.00 0.00 11.5 -2.91 -0.12 60.0 141.96 30.67 71.5 118.74 30.5542 8158 -1.02 -30.07 2646.6 -2.92 -27.84 933.0 125.48 421.46 11737.3 8.60 363.5544 0 0.00 0.00 74.2 0.00 0.00 6.1 90.32 1.97 80.3 6.83 1.9746 0 0.00 0.00 15.0 0.00 0.00 1.0 354.97 1.23 16.0 21.42 1.2353 8158 0.65 19.10 169.5 -2.91 -1.78 10.8 114.01 4.44 8338.0 0.72 21.7554 0 0.00 0.00 2360.7 -2.91 -24.74 99.2 114.01 40.72 2459.9 1.80 15.9855 0 0.00 0.00 15.0 0.00 0.00 35.9 140.14 18.10 50.9 98.77 18.1056 0 0.00 0.00 15.0 0.00 0.00 0.0 0.00 0.00 15.0 0.00 0.0057 0 0.00 0.00 19.7 -2.91 -0.21 17.1 128.29 7.89 36.8 57.94 7.6858 8158 -1.02 -30.07 2568.9 -2.91 -26.93 384.4 114.01 157.77 11111.0 2.52 100.7859 7488 3.82 102.93 1.6 0.00 0.00 0.0 0.00 0.00 7490.1 3.82 102.9360 8158 3.14 92.20 88.0 -2.91 -0.92 0.7 114.01 0.29 8246.4 3.08 91.5761 0 0.00 0.00 22.1 -2.91 -0.23 115.5 141.96 59.05 137.6 118.74 58.8262 0 0.00 0.00 32.1 0.00 0.00 -0.1 0.00 0.00 31.9 0.00 0.0063 0 0.00 0.00 22.1 0.00 0.00 1.1 141.96 0.57 23.2 6.82 0.5764 0 0.00 0.00 51.3 7.36 1.36 71.7 257.31 66.43 123.0 153.08 67.7967 0 0.00 0.00 72.2 11.15 2.90 4.0 33.65 0.48 76.2 12.32 3.3868 0 0.00 0.00 97.9 -0.53 -0.19 24.5 238.55 21.02 122.4 47.28 20.8370 751 3.82 10.33 0.2 0.00 0.00 0.0 0.00 0.00 751.6 3.82 10.3371 10062 -2.71 -98.29 1551.3 -2.79 -15.58 76.1 256.55 70.31 11689.8 -1.04 -43.5672 0 0.00 0.00 18.8 3.82 0.26 1.6 3.82 0.02 20.4 3.82 0.2873 0 0.00 0.00 1.1 -2.91 -0.01 5.9 141.96 2.99 7.0 118.74 2.9875 8158 3.14 92.20 88.0 -2.91 -0.92 0.0 0.00 0.00 8245.7 3.07 91.2876 10062 -2.71 -98.29 1551.3 -2.79 -15.58 145.7 240.42 126.07 11759.4 0.29 12.2078 0 0.00 0.00 1.1 -2.91 -0.01 0.1 141.96 0.03 1.2 4.03 0.0280 0 0.00 0.00 108.0 -3.05 -1.19 265.8 33.65 32.20 373.9 23.04 31.0181 0 0.00 0.00 9.7 0.00 0.00 0.0 0.00 0.00 9.7 0.00 0.0082 0 0.00 0.00 9.7 0.00 0.00 0.2 375.68 0.29 9.9 7.99 0.2985 0 0.00 0.00 32.1 0.00 0.00 81.5 223.45 65.60 113.6 160.38 65.6086 0 0.00 0.00 108.0 -3.05 -1.19 54.3 33.65 6.58 162.3 9.22 5.3987 0 0.00 0.00 108.0 -3.05 -1.19 26.2 33.65 3.18 134.2 4.11 1.9990 0 0.00 0.00 1.5 0.00 0.00 2.9 439.90 4.60 4.5 287.29 4.6091 0 0.00 0.00 15.0 0.00 0.00 10.9 311.93 12.29 26.0 131.44 12.2992 0 0.00 0.00 3.8 -2.91 -0.04 20.0 141.96 10.22 23.8 118.74 10.1893 0 0.00 0.00 3.8 -2.91 -0.04 0.2 141.96 0.09 4.0 3.16 0.0594 0 0.00 0.00 873.6 -2.91 -9.16 4.5 114.01 1.86 878.2 -2.31 -7.3095 0 0.00 0.00 35.8 -3.05 -0.39 3.3 33.65 0.40 39.1 0.04 0.0196 8158 -1.02 -30.07 2530.5 -2.91 -26.52 139.9 114.01 57.43 10828.1 0.02 0.8499 0 0.00 0.00 33.6 -2.91 -0.35 0.0 114.01 0.00 33.6 -2.91 -0.35

100 8240 3.82 113.26 1.8 0.00 0.00 0.0 0.00 0.00 8241.7 3.82 113.26

266


Table B.27: Component results from the exergoeconomic analysis of HTC-1.00-s.

No type ED EL yD yD* EF EP cF cP r f CD Z CD+Z[kW] [kW] [%] [%] [kW] [kW] [%] [€/GJ] [€/GJ] [-] [-] [€/h] [€/h] [€/h]

1 slurry-mix 18.6 0.0 0.4% 1.0% 38.6 20.1 51.9% 74.8 184.8 147% 37% 5.00 2.94 7.942 slurry pump 2.3 0.0 0.0% 0.1% 3.3 0.9 28.6% 28.7 656.1 2189% 89% 0.24 1.87 2.123 slurry preheater HX 9.0 0.0 0.2% 0.5% 43.7 34.8 79.5% 171.9 258.9 51% 49% 5.54 5.34 10.884 slurry preheater HX 21.6 0.0 0.4% 1.2% 62.7 41.2 65.6% 155.6 256.1 65% 19% 12.08 2.81 14.905 steam / slurry mix 26.6 0.0 0.5% 1.5% 127.4 100.8 79.1% 150.6 190.4 26% 0% 14.42 0.00 14.426 slurry pump 7.0 0.0 0.1% 0.4% 10.8 3.8 35.2% 28.7 433.4 1412% 87% 0.72 4.80 5.527 steam / slurry mix 34.2 0.0 0.7% 1.9% 247.7 213.5 86.2% 133.4 154.8 16% 0% 16.42 0.00 16.428 slurry pump 7.2 0.0 0.1% 0.4% 12.1 4.8 40.0% 28.7 397.2 1286% 88% 0.75 5.66 6.409 steam / slurry mix 10.0 0.0 0.2% 0.6% 127.4 117.5 92.2% 161.8 175.5 8% 0% 5.81 0.00 5.8110 slurry pump 15.0 0.0 0.3% 0.8% 25.9 10.9 42.1% 28.7 361.7 1162% 88% 1.55 11.51 13.0511 heat loss 7.4 0.0 0.1% 0.4% 7.412 air preheater 2.6 0.0 0.1% 0.1% 5.3 2.7 51.4% 171.9 504.9 194% 51% 1.59 1.66 3.2513 steam / slurry mix 16.3 0.0 0.3% 0.9% 248.7 232.4 93.4% 43.2 46.3 7% 0% 2.54 0.00 2.5414 HTC reactor 505.0 0.0 9.9% 27.9% 3264.7 2759.7 84.5% 5.0 12.7 151% 88% 9.18 66.74 75.9215 evaporator (offgas) 9.5 0.0 0.2% 0.5% 102.2 92.7 90.7% 45.5 58.7 29% 64% 1.56 2.83 4.3916 heat loss 9.1 0.0 0.2% 0.5% 9.117 flash 1 8.6 0.0 0.2% 0.5% 124.6 116.0 93.1% 148.1 177.4 20% 62% 4.59 7.64 12.2318 flash 2 0.0 0.0 0.0% 0.0% 222.4 222.4 100.0% 138.1 144.8 5% 100% 0.00 5.35 5.3519 flash 3 20.2 0.0 0.4% 1.1% 192.4 172.2 89.5% 138.1 159.5 15% 24% 10.03 3.22 13.2520 flash 4 43.6 0.0 0.9% 2.4% 212.0 168.4 79.5% 138.6 178.9 29% 11% 21.73 2.68 24.4121 mechanical dewatering 43.6 0.0 0.9% 2.4% 28.7 89% 4.50 36.74 41.2322 drier 64.1 0.0 1.3% 3.5% 294.5 26% 67.92 24.26 92.1824 condensate cooler 9.8 0.0 0.2% 0.5% 149.0 34% 5.24 2.73 7.9825 boiler 438.4 0.0 8.6% 24.2% 693.6 254.8 36.7% 15.2 49.9 229% 25% 23.92 7.95 31.8726 FWPH 9.4 0.0 0.2% 0.5% 49.3 40.0 81.0% 45.5 70.5 55% 57% 1.53 2.06 3.5927 pump 0.9 0.0 0.0% 0.1% 1.9 0.9 50.2% 28.7 366.4 1178% 92% 0.10 1.05 1.1428 condenser 0.0 0.0 0.0% 0.0% 0.0 1.9829 mix 11.4 0.0 0.2% 0.6% 905.0 893.6 98.7% 136.3 138.0 1% 0% 5.61 0.00 5.6131 mix 0.0 0.0 0.0% 0.0% 387.6 387.6 100.0% 138.4 138.5 0% 0% 0.02 0.00 0.0232 drier HX 60.4 0.0 1.2% 3.3% 124.5 64.1 51.5% 171.9 355.4 107% 10% 37.39 4.12 41.5133 waste water cooler 8.2 0.0 0.2% 0.5% 138.5 29% 4.10 1.69 5.7936 FWPH 10.2 0.0 0.2% 0.6% 19.9 9.7 48.9% 171.9 377.5 120% 12% 6.28 0.90 7.1837 biocoal cooler 0.1 0.0 0.0% 0.0% 138.5 83% 0.06 0.33 0.3938 drier fan 11.9 0.0 0.2% 0.7% 34.5 22.6 65.5% 28.7 45.3 58% 9% 1.23 0.12 1.3539 drier gas mix 51.7 0.0 1.0% 2.9% 178.9 127.2 71.1% 143.1 201.3 41% 0% 26.64 0.00 26.6440 pump 0.2 0.0 0.0% 0.0% 0.3 0.1 45.1% 28.7 441.4 1440% 92% 0.02 0.18 0.2041 FWPH (offgas) 16.1 0.0 0.3% 0.9% 35.9 19.8 55.2% 45.5 96.0 111% 27% 2.63 0.96 3.6042 offgas flash 37.4 0.0 0.7% 2.1% 102.2 64.0 62.6% 20.9 41.4 99% 41% 2.81 1.99 4.8043 boiler fuel valve 3.3 0.0 0.1% 0.2% 67.6 64.3 95.1% 41.7 43.8 5% 0% 0.50 0.00 0.5044 heat loss biocoal storage 0.6 0.0 0.0% 0.0% 0.6 138.5 0% 0.31 0.00 0.3145 condensate collector 0.6 0.0 0.0% 0.0% 125.0 124.4 99.5% 18.2 18.2 0% 0% 0.04 0.00 0.0446 pellet press 74.2 0.0 1.5% 4.1% 70.2 28.7 69% 7.66 16.67 24.3347 misc. el. consumption 185.6 0.0 3.6% 10.2% 185.6 28.7 0% 19.15 0.00 19.15

total plant 1811.7 944.2 35.4% 100% 5113.3 2357.8 46.1%

B.4.7 Alternative flowsheet designs HTC-3.20 to HTC-3.80

In HTC-3.20 with SSD at atmospheric pressure (Figure B.2), the drier is heated withsteam at 6 bar. The condensate (63) and the exhaust steam of the drier (68) are led tothe flash tank at 1 bar (K20). Steam at 1 bar is recompressed and used as fluidizationsteam (48) for the drier. Furthermore, 1 bar steam is compressed to 6 bar (K38) to coverthe demand of 6 bar steam for preheating the biomass slurry and supplying the drier.1

HTC-3.40 is similar to HTC-3.30, but the SSD is operated at 4 bar, and the drier exhauststeam is fed to the flash tank (K19). An additional flash tank (K38) is introduced intothe plant design to provide steam at 7 bar for the drier fluidization. Steam at 15 bar isutilized to heat the drier.In HTC-3.50 (Figure B.5), the 1 bar steam released by flash evaporation during de-pressurization (K34) is led to the 1 bar flash tank.

1Like in the base case, it is assumed that 30% of the dissolved organics TOMres are lost as VOC emissionsin the drier. While in reality they would leave the drier with the exhaust steam and end up in thecondensate (36), they are purged directly from the drier (99) in the simulation.

267


In HTC-3.80 without heat transfer from the biocoal to the biomass slurry (Figure B.6),the biocoal slurry is de-pressurized to 1 bar in one step (K20). Steam at 1 bar is used toprovide thermal energy to the drier, the remainder (78) is discharged to the condenser.

condensate

reactor

filter press

76

50

14

48

48

62

47

19 20 21 22 23

6

15

13

9654

53

1639

37 4

10

7

5 67

8011 86

12

87

82

81

5692

41

97

28

61

85

64

78

70

6029

90

7873

95

95

93

93

9

32

34

46

91

55

8340

8

57

44

36

58 6863

99

30

42

K2 K3 K4 K5

K28

K24

K45

K6 K7 K8 K9 K10 K11

K16 K15

K29

K19

K31

K21 K37K22

K20

K18

K41

K40

K36

K27

K12

K39

K38

K38

K26

K25

K42

K43

K13

K14

K17

17 71

air

SSD

A

A

BC

C

D

D

E

B

biomass

VOCloss

fluegas

27

1 K1

K33

3

2W1 W2 W4

W7

W5

W11

W14

W8

W3

wastewater

aeration

discharge

94

W14

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

biocoal

pelletpress

70

59

100

K44K46

E

W9

ash72


biocoal

condensate

reactor

pelletpress

incinerator

filter press

76

50

14

48

48

62

47

19 20 21 22 23

6

15

13

9654

53

1639

37 4

10

5

678011 86

65

12

82

8151

84

41

97

28

61

85

64

78

5960

29

90

7873

95

95

9

32

34

8340

8

57

44

36

58 6863

99

30

42

K2 K3 K4 K5

K28

K24

K45

K6 K7 K8 K9 K10 K11

K16 K15

K29

K19

K31

K21 K37K46

K22

K20

K18

K41

K40

K12

K39

K32

K27

K38

K42

K13 K25

K26K14

K17

17 71

air

SSD

A

A

C

C

D

D

biomass

VOCloss

fluegas

27

1 K1

K33

3

2 W1 W2 W4

W9W7

W5

W3

W11

W14

W8

wastewater

aeration

discharge

94

W15

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

K44


268


condensate

reactor

filter press

76

50

14

48

48

62

47

19 20 21 22 23

6

15

65

43

13

9654

53

1639

37 4

10

7

5 67

8011 86

12

87

82

81

5692

41

97

28

6164

64

78

70

6029

90

7873

95

95

93

93

9

32

34

46

91

55

83

40

8

85

57

44

36

58 6863

99

30

42

K2 K3 K4 K5

K28

K24

K45

K6 K7 K8 K9 K10 K11

K16 K15

K29

K19

K31

K21 K37K22

K20

K18

K41

K40

K36

K27

K12K39

K38

K26

K25

K42

K43

K13

K14

K17

17 71

air

SSD

A

A

BC

C

D

F

F

D

E

B

biomass

VOCloss

fluegas

27

1 K1

K33

3

2W1 W2 W4

W7

W5

W11

W3

wastewater

aeration

discharge

94

W14

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

biocoal

pelletpress

70

59

100

K44K46

E

W9

ash72


condensate

reactor

flue gas

air

76

50

14

47

19 20 21 22 23

6

15

13

96

54

53

48

48

1639

37 4

10

7

5 67

8011 86

12

87

82

81

56

62

92

41

97

61

28

63

78

8564

70

60

29

90

7873

63

95

95

93

93

9

32

34

46

91

55

8340

8

57

44

36

5868+99

30 74

42

K2 K3 K4 K5

K28

K24

K45

K6 K7 K8 K9 K10 K11

K16

K15

K29

K19

K31

K21

K37

K22

K20 K34

K18

K41

K40

K36

K27

K12

K39

K38K32

K26

K25

K42

K43

K13

K14

K17

17 71

air

drier

A

A

BC

C

D

F

F

E

D

B

biomass

filterpress

27

1 K1

K33

3

2W1 W2 W4

W7

W5

W11

W8

W3

wastewater

aeration

discharge

94

W14

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

biocoal

pelletpress

70

59

100

K44K46

E

W9


condensate

reactor

flue gas

filter press

air

6

96

54

53

4

10

7

6780

56

62

78 61

28

6364

70

60

29

63

95

95

9

32

55

44

3668+99

3042

K2

K28

K24

K45

K16

K21 K37

K22

K20

K27

K38K32

K25

K42

K13

K14air

drier

A

A

D

E

D

biomass

fluegas

27

1 K1

K33

3

2 W1

W7

W8

W3

wastewater

aeration

discharge

94

W14

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

biocoal

pelletpress

70

59

100

K44K46

E

W9

ash72


269


B.4.8 Summarized results from HTC-3.01 to HTC-3.90

Table B.28: Results from HTC-3.20 to HTC-3.90. Biomass input is 9509 kWHHV PGW-70 in all cases.

HTC- 3.20 3.30 3.40 3.50 3.60 3.70 3.80 3.90

biocoal yield [kW] 7590 7925 7013 7284 7369 7921 4572 7692

electricity consumption [kW] 545 690 413 484 403 429 284

carbon yield [–] 81.6% 85.2% 75.4% 78.3% 79.2% 85.1% 49.1% 82.6%

biocoal combusted [–] 4.2% 0.0% 11.5% 7.9% 6.9% 0.0% 42.3% 2.9%

HTC reactor [M€] 1.78 1.76 1.76 1.64 1.57 1.93 1.77

slurry pumps [M€] 0.50 0.50 0.47 0.86 0.33 0.33 0.47

flash tanks [M€] 0.54 0.54 0.51 0.09 0.23 0.13 0.53

filter press [M€] 0.75 0.74 1.53 1.32 0.74 0.74 0.74

biocoal drier (incl. HX) [M€] 0.66 0.96 0.57 0.57 0.66 0.78 0.78

pellet press [M€] 0.34 0.34 0.34 0.34 0.34 0.34 0.34

boiler [M€] 0.21 0.21 0.20 0.23 0.28 0.34 0.23

heat exchangers [M€] 0.46 0.41 0.52 0.26 0.60 0.00 0.51

coolers [M€] 0.44 0.42 0.29 0.33 0.02 0.61 0.34

steam compressor [M€] 0.42 0.79 0.00 0.00 0.00 0.00

waste water treatment [M€] 0.34 0.32 0.35 0.32 0.28 0.40 0.34

digester 0.51

other [M€] 0.54 0.59 0.45 0.39 0.43 0.45 0.45

total CBM [M€] 6.98 7.57 7.00 6.35 5.48 6.05 7.02

TCI [M€] 10.73 11.53 10.75 9.92 8.75 9.50 10.74

carrying charges [€GJ] 7.89 8.11 8.23 7.51 6.16 11.58 7.78

labour [€/GJ] 5.33 5.11 5.55 5.49 5.11 8.85 5.26

electricity [€/GJ] 2.06 2.50 1.62 1.88 1.46 2.69 1.06

O&M, material [€/GJ] 1.39 1.39 1.78 1.77 1.12 2.03 1.36

other operating cost [€/GJ] 0.02 0.02 0.02 0.02 0.02 0.02 0.02

biomass incl. transport [€/GJ] -3.60 -3.45 -3.75 -3.71 -3.45 -5.97 -3.55

total biocoal cost [€/GJ] 13.09 13.68 13.46 12.96 10.42 19.20 11.94

270


Tabl

eB

.29:

Res

ults

from

HT

C-3

.01

toH

TC

-3.1

3.B

iom

ass

inpu

tis

9509

kWH

HV

PG

W-7

0in

allc

ases

.

HT

C-

3.01

3.02

3.03

3.04

3.05

3.06

3.07

3.08

-p1

3.08

-p2

3.09

3.10

3.11

bioc

oaly

ield

[kW

]71

5969

6970

3972

0672

9568

4072

8173

0573

0573

1973

3273

37

elec

tric

ity

cons

umpt

ion

[kW

]40

742

042

140

341

641

041

638

149

240

740

741

6

carb

onyi

eld

[–]

76.8

%75

.0%

75.6

%77

.4%

78.4

%73

.5%

78.3

%78

.5%

78.5

%78

.6%

78.8

%78

.8%

bioc

oalc

ombu

sted

[–]

10.5

%10

.9%

11.2

%9.

0%7.

9%14

.0%

7.8%

7.7%

7.7%

7.6%

7.5%

7.4%

HT

Cre

acto

r[M

€]

1.30

3.55

1.12

2.96

1.84

1.78

1.77

1.06

1.06

1.79

1.79

1.84

slur

rypu

mps

[M€

]0.

410.

530.

530.

410.

510.

460.

480.

291.

010.

450.

450.

50

flash

tank

s[M

€]

0.52

0.61

0.62

0.51

0.53

0.53

0.53

0.34

0.34

0.53

0.53

0.53

filte

rpr

ess

[M€

]0.

750.

720.

730.

740.

740.

750.

730.

450.

450.

740.

740.

74

bioc

oald

rier

(inc

l.H

X)

[M€

]0.

800.

740.

780.

780.

781.

010.

570.

780.

780.

780.

780.

78

pelle

tpr

ess

[M€

]0.

350.

320.

340.

340.

340.

340.

340.

340.

340.

340.

340.

34

boile

r[M

€]

0.23

0.31

0.32

0.21

0.25

0.33

0.22

0.20

0.20

0.20

0.19

0.22

heat

exch

ange

rs[M

€]

0.51

0.56

0.56

0.51

0.44

0.47

0.49

0.26

0.26

0.38

0.34

0.33

cool

ers

[M€

]0.

340.

370.

350.

350.

330.

350.

390.

330.

330.

320.

310.

31

was

tew

ater

trea

tmen

t[M

€]

0.34

0.36

0.35

0.34

0.32

0.34

0.34

0.34

0.34

0.31

0.31

0.31

othe

r[M

€]

0.46

0.45

0.46

0.45

0.45

0.46

0.45

0.43

0.43

0.45

0.45

0.45

tota

lC

BM

[M€

]6.

018.

526.

157.

606.

526.

826.

304.

815.

546.

286.

226.

35

TC

I[M

€]

9.44

12.7

39.

6311

.52

10.1

110

.50

9.82

7.86

8.84

9.80

9.72

9.89

carr

ying

char

ges

[€G

J]7.

3510

.18

7.63

8.91

7.73

8.56

7.52

6.00

6.75

7.46

7.39

7.51

labo

ur[€

/GJ]

5.65

5.81

5.75

5.61

5.55

5.91

5.56

5.54

5.54

5.53

5.52

5.51

elec

tric

ity

[€/G

J]1.

631.

731.

711.

601.

641.

721.

641.

501.

931.

591.

591.

62

O&

M,m

ater

ial

[€/G

J]1.

331.

681.

411.

491.

401.

511.

360.

991.

381.

341.

331.

37

othe

rop

erat

ing

cost

[€/G

J]0.

020.

020.

020.

020.

020.

020.

020.

020.

020.

020.

020.

02

biom

ass

incl

.tr

ansp

ort

[€/G

J]-3

.81

-3.9

2-3

.88

-3.7

9-3

.74

-3.9

9-3

.75

-3.7

4-3

.74

-3.7

3-3

.72

-3.7

2

tota

lb

ioco

alco

st[€

/GJ]

12.1

715

.50

12.6

513

.86

12.5

813

.73

12.3

410

.31

11.8

912

.22

12.1

412

.32

271


B.4.9 Simulation data from HTC-3.30 and HTC-3.60


stream no. 1 2 3 4 5 6 8 9 10 11 12 13 14 15 16 17 19 20T [°C] 15.0 90.0 56.7 187.9 314.1 219.0 159.1 99.9 203.7 200.0 190.1 188.3 159.1 206.7 206.7 56.9 138.1 138.5p [bar] 1.0 1.0 1.0 29.0 3.0 27.0 6.0 1.0 29.0 25.0 12.6 12.0 6.0 18.0 18.0 4.0 4.0 10.0m [kg/h] 6666.7 6969.7 13636.4 16030.0 197.8 15663.2 13533.8 1470.4 16383.6 720.4 380.0 997.9 1886.9 474.6 15188.6 13636.4 14557.5 14557.5H [kW] -23536 -29274 -52809 -60887 -468 -59717 -51878 -6235 -62061 -2583 -1602 -3635 -6906 -1700 -58017 -52806 -55570 -55562O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 0.00 0.00 17.41 181.88 17.43 0.00 0.90 17.42 182.77 0.00 1.27 0.04 16.12 1.30 0.00 0.02 0.02CO [kg/h] 0.00 0.00 0.00 3.38 8.12 3.42 0.00 0.33 3.39 8.41 0.00 0.83 0.13 2.48 0.94 0.00 0.06 0.06H2 [kg/h] 0.00 0.00 0.00 0.04 0.71 0.04 0.00 0.00 0.04 0.71 0.00 0.00 0.00 0.04 0.00 0.00 0.00 0.00H2O [kg/h] 4666.67 6666.67 11333.33 13691.16 6.33 13977.79 11885.68 1461.56 14043.16 524.44 380.00 989.40 1876.10 452.66 13525.13 11333.33 12249.11 12249.11CH4 [kg/h] 0.00 0.00 0.00 0.35 0.30 0.37 0.01 0.06 0.36 0.34 0.00 0.14 0.06 0.18 0.19 0.00 0.03 0.03N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/h] 0.00 0.00 0.00 2.22 0.49 2.84 0.42 0.70 2.39 0.74 0.00 1.02 1.19 0.62 2.22 0.00 0.58 0.58C2H4O2 [kg/h] 0.00 114.29 114.29 125.64 0.01 199.74 187.16 6.31 126.96 2.75 0.00 4.82 8.65 2.30 197.44 114.29 118.52 118.52CH2O2 [kg/h] 0.00 7.96 7.96 8.93 0.00 14.23 13.18 0.57 9.06 0.22 0.00 0.41 0.76 0.19 14.04 7.96 8.33 8.33biocoal DM [kg/h] 0.00 0.00 0.00 0.00 0.00 1139.82 1139.82 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1139.82 0.00 0.00 0.00biomass DM [kg/h] 2000.00 0.00 2000.00 2000.00 0.00 0.00 0.00 0.00 2000.00 0.00 0.00 0.00 0.00 0.00 0.00 2000.00 2000.00 2000.00TOMres [kg/h] 0.00 95.19 95.19 95.19 0.00 161.89 161.89 0.00 95.19 0.00 0.00 0.00 0.00 0.00 161.89 95.19 95.19 95.19ash, diss. [kg/h] 0.00 85.63 85.63 85.63 0.00 145.63 145.63 0.00 85.63 0.00 0.00 0.00 0.00 0.00 145.63 85.63 85.63 85.63



stream no. 78 80 81 82 83 84 85 86 90 94 95 96 97 99T [°C] 43.1 218.5 15.0 15.3 188.0 733.2 364.1 171.6 90.1 40.0 60.0 100.1 100.0 100.0p [bar] 1.0 25.0 1.0 12.9 12.0 31.0 6.0 25.0 1.0 1.0 3.0 1.0 1.0 1.0m [kg/h] 2.3 720.4 380.0 380.0 380.0 353.6 850.0 720.4 73.5 4089.0 522.6 12993.1 945.6 3.3H [kW] -10 -2373 -1679 -1679 -1391 -1174 -3010 -2661 0 -17399 -2273 -50405 -3952 -3O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 17.01 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 182.77 0.00 0.00 0.00 0.00 0.01 182.77 0.00 0.00 0.89 0.00 0.01 0.00CO [kg/h] 0.00 8.41 0.00 0.00 0.00 0.01 0.03 8.41 0.00 0.00 0.30 0.00 0.03 0.00H2 [kg/h] 0.00 0.71 0.00 0.00 0.00 0.00 0.00 0.71 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 2.26 524.44 380.00 380.00 380.00 351.99 846.04 524.44 0.47 3911.23 518.11 11337.78 941.19 0.00CH4 [kg/h] 0.00 0.34 0.00 0.00 0.00 0.01 0.01 0.34 0.00 0.00 0.05 0.00 0.02 0.00N2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 56.01 0.00 0.00 0.00 0.00 0.00H2S [kg/h] 0.00 0.74 0.00 0.00 0.00 0.17 0.40 0.74 0.00 0.00 0.25 0.09 0.45 0.00C2H4O2 [kg/h] 0.01 2.75 0.00 0.00 0.00 1.33 3.19 2.75 0.00 67.05 2.74 194.38 3.55 0.00CH2O2 [kg/h] 0.00 0.22 0.00 0.00 0.00 0.13 0.31 0.22 0.00 4.67 0.22 13.54 0.34 0.00biocoal DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1139.82 0.00 0.00biomass DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00TOMres [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 55.85 0.00 161.89 0.00 3.26ash, diss. [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 50.24 0.00 145.63 0.00 0.00

[kWel] heat losses [kWth]W1 4.3 slurry pump (K2) biomass slurry (K11) 19.1W2 9.8 slurry pump (K6) biocoal slurry (K16) 20.3W3 147.6 compressor (K27) boiler (K25) 0.0W4 10.9 slurry pump (K8) HTC reactor (K14) 8.1W5 24.8 slurry pump (K10) drier (K22) 21.0W7 13.1 filter press (K21) cooler duty [kWth]W8 18.5 compressor (K38) reactor offgas (K42) 96.3W9 64.3 pellet press (K46) condenser (K28) 184.2W10 689.8 total plant condensate (K24) 18.7W11 0.3 pump (K40) waste water (K33) 224.3W13 1.0 cooler fans filter press (K21) 139.3W14 141.4 compressor (K32)W15 253.7 waste water aeration

272





stream no. 68 70 72 73 75 80 85 86 90 94 95 99 100T [°C] 46.8 15.0 800.0 99.8 15.0 218.7 89.8 70.0 89.8 40.0 59.0 45.1 15.0p [bar] 1.0 1.0 1.0 1.0 1.0 28.0 1.1 28.0 1.0 1.0 3.0 1.0 1.0m [kg/h] 12831.9 88.1 5.6 28.4 1282.5 419.3 11300.0 419.3 920.9 3974.7 252.9 2.7 1282.5H [kW] -2593 -94 0 -104 -1362 -1322 -60 -1511 -5 -16897 -1101 -2 -1363O2 [kg/h] 2682.23 0.00 0.00 0.00 0.00 0.00 2615.56 0.00 213.16 0.00 0.00 0.00 0.00CO2 [kg/h] 340.18 0.00 0.00 0.51 0.00 156.17 0.00 156.17 0.00 0.01 0.45 0.00 0.00CO [kg/h] 0.01 0.00 0.00 0.07 0.00 4.88 0.00 4.88 0.00 0.01 0.10 0.00 0.00H2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.65 0.00 0.65 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 485.69 8.66 0.00 27.74 126.05 256.04 71.54 256.04 5.83 3800.05 250.97 0.00 126.05CH4 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.12 0.00 0.12 0.00 0.00 0.01 0.00 0.00N2 [kg/h] 9316.41 0.00 0.00 0.00 0.00 0.00 8612.91 0.00 701.92 0.00 0.00 0.00 0.00H2S [kg/h] 0.02 0.00 0.00 0.02 0.00 0.15 0.00 0.15 0.00 0.04 0.04 0.01 0.00C2H4O2 [kg/h] 6.46 0.00 0.00 0.10 2.29 1.23 0.00 1.23 0.00 60.50 1.23 0.00 0.00CH2O2 [kg/h] 0.44 0.00 0.00 0.01 0.16 0.10 0.00 0.10 0.00 4.21 0.10 0.00 0.00biocoal DM [kg/h] 0.00 79.42 0.00 0.00 1139.81 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1156.41biomass DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00TOMres [kg/h] 0.00 0.00 0.00 0.00 6.19 0.00 0.00 0.00 0.00 57.85 0.00 2.65 0.00ash, diss. [kg/h] 0.00 0.00 4.12 0.00 7.96 0.00 0.00 0.00 0.00 52.04 0.00 0.00 0.00

[kWel] heat losses [kWth]W1 132.4 slurry pump (K2) biomass slurry (K11) 0.0W3 1.8 pump (K27) biocoal slurry (K16) 18.0W4 1.9 pump (K40) filter press (K21) 33.1W7 14.0 filter press (K21) boiler (K25) 17.7W8 14.4 drier fan (K38) HTC reactor (K14) 8.1W9 64.1 pellet press (K46) drier (K22) 11.4W10 483.6 total plant cooler duty [kWth]W13 3.7 cooler fans reactor offgas (K42) 0.0W14 251.3 waste water aeration condenser (K28) 157.0

condensate (K24) 152.2waste water (K33) 261.7

273


B.4

.10

Cost

data

from

HT

C-3.

30an

dH

TC-

3.60

Tabl

eB

.32:

HT

C-3

.30-

seq

uipm

ent

list

wit

hin

vest

men

tco

sts.

com

pone

ntno

.un

itX

sim

f dn

Xd/

nC

BM

/n [k€

]

CB

M

[k€

]

spec

ifica

tion

sco

stfu

ncti

on

HT

Cre

acto

rK

14[m

3]

84.7

712

0%2

50.8

688

217

63SS

-cla

d

p=p n

om+

5ba

r,h/

d=5

HT

C-1

,HT

C-2

,HT

C-3

slur

rym

ixin

g

tank

K1

[m3]

7.41

120%

18.

8966

66re

side

nce

tim

e:30

min

,

Vto

tal/

Va

cti

ve=

1.1,

l/d=

2,

SS-c

lad

TN

K-1

,T

NK

-2

flash

tank

1K

17[m

3]

12.5

912

0%1

15.1

121

721

7re

side

nce

tim

e:20

min

(liq

uid)

,liq

uid

fillin

g50

%,

l/d=

2,SS

-cla

d

TN

K-1

,T

NK

-2

flash

tank

2K

18[m

3]

11.7

512

0%1

14.1

015

415

4se

efla

shta

nk1

TN

K-1

,T

NK

-2

flash

tank

3K

19[m

3]

10.4

912

0%1

12.5

993

93se

efla

shta

nk1

TN

K-1

,T

NK

-2

flash

tank

4K

20[m

3]

9.42

120%

111

.30

7676

see

flash

tank

1T

NK

-1,

TN

K-2

slur

rypu

mp

1K

2[m

3/s

]3.

7912

0%2

2.27

2754

pum

ping

in4

stag

es,t

otal

Δp=

31ba

r

SLP

-11)

slur

rypu

mp

2K

6[m

3/s

]8.

5712

0%2

5.14

4691

see

slur

rypu

mp

1se

esl

urry

pum

p1

slur

rypu

mp

3K

8[m

3/s

]9.

4712

0%2

5.68

5410

8se

esl

urry

pum

p1

see

slur

rypu

mp

1

slur

rypu

mp

4K

10[m

3/s

]21

.61

120%

212

.96

117

234

see

slur

rypu

mp

1se

esl

urry

pum

p1

filte

rpr

ess

K21

[m3/h

]14

.18

120%

117

.01

741

741

high

pres

sure

,SS

FP

-1,F

P-2

inci

nera

tor

K25

[kW

]57

.61

120%

169

.13

100

100

OG

I-1

gas

boile

r(f

orst

art-

up)

[kg/

h]0.

001

3200

106

106

30ba

rG

B-1

slur

ryH

XK

3[m

2]

13.5

811

5%1

15.6

219

619

6sp

iral

plat

e,SS

/SS

HX

S-1

slur

ryH

XK

4[m

2]

1.49

115%

11.

7164

64sp

iral

plat

e,SS

/SS

HX

S-1

offga

sH

XK

15[m

2]

14.3

311

5%1

16.4

882

82SS

/C

SH

X-1

offga

sH

XK

41[m

2]

1.82

115%

12.

0932

32SS

/C

SH

X-1

air

preh

eate

rK

12[m

2]

0.23

115%

10.

269

9SS

/C

SH

X-1

recu

pera

tor

K26

[m2]

1.31

115%

11.

5126

26SS

/C

SH

X-1

274


com

pone

ntno

.un

itX

sim

f dn

Xd/

nC

BM

/n [k€

]

CB

M

[k€

]

spec

ifica

tion

sco

stfu

ncti

on

coal

cool

erK

37[m

2]

1.36

115%

11.

5735

35ai

r-co

oled

,SS

AC

O-1

coal

cool

erK

96[m

2]

3.20

115%

13.

6857

57ai

r-co

oled

,SS

AC

O-1

cond

ense

rK

28[m

2]

6.11

115%

17.

0284

84ai

r-co

oled

,SS

AC

O-1

cond

ensa

teco

oler

K24

[m2]

6.73

115%

17.

7489

89ai

r-co

oled

,SS

AC

O-1

purg

eco

oler

K33

[m2]

9.35

115%

110

.75

106

106

air-

cool

ed,S

SA

CO

-1

offga

sco

oler

K42

[m2]

2.57

115%

12.

9650

50ai

r-co

oled

,SS

AC

O-1

pum

pK

40[k

W]

0.28

110%

10.

315

5C

SP

-1

stea

mco

mpr

.K

38[k

W]

15.7

011

0%1

17.2

662

62SS

CM

P-1

mot

orK

38[k

W]

18.5

211

0%1

20.3

717

17E

M-1

stea

mco

mpr

.K

32[k

W]

126.

8011

0%1

139.

4735

635

6SS

CM

P-1

mot

orK

32[k

W]

141.

4011

0%1

155.

5476

76E

M-1

stea

mco

mpr

.K

27[k

W]

132.

5411

0%1

145.

7937

037

0SS

CM

P-1

mot

orK

27[k

W]

147.

6311

0%1

162.

4078

78E

M-1

biom

ass

scre

wco

nvey

or[k

g/h_

ar]

6667

120%

180

0020

20l=

45m

SCC

-1

bioc

oalS

SDK

22[k

g ev/h

]63

312

0%1

760

960

960

SSD

-1

was

tew

ater

trea

tmen

tpl

ant

[m3/h

]5.

5812

0%1

6.69

317

317

WW

T-1

pelle

tpr

ess

K46

[kg/

h]12

8612

0%1

1543

338

338

x(P

P-1

,PP

-2)+

PC

-1+

PSS

-1

pelle

tsco

alst

orag

e&

hand

ling

[kg/

h]12

8612

0%1

1543

166

166

see

case

HT

C-3

.00-

sP

S-1

biom

ass

sizi

ng&

sort

ing

[kg/

h]66

6712

0%1

8000

162

162

shre

dder

,dru

msi

eve,

mag

neti

cse

para

tor,

air

clas

sifie

r

BSS

-1+

BSS

-2+

BSS

-3

tota

lC

BM

7571

1)C

BM

isca

lcul

ated

for

tota

lΔp=

31ba

r,th

enas

sign

edto

the

stag

es1-

4ba

sed

onth

eirΔ

ppe

rst

age

275


Tabl

eB

.33:

HT

C-3

.60-

seq

uipm

ent

list

wit

hin

vest

men

tco

sts.

com

pone

ntno

.un

itX

sim

f dn

Xd/

nC

BM

/

[k€

]

CB

M

[k€

]

spec

ifica

tion

sco

stfu

ncti

on

HT

Cre

acto

rK

14[m

3]

73.5

012

0%2

44.1

082

116

41SS

-cla

d

p=p n

om+

5ba

r,h/

d=5

HT

C-1

,HT

C-2

,HT

C-3

depr

essu

riza

tion

tank

K17

[m3]

1.37

120%

11.

6426

26re

side

nce

tim

e:20

min

(slu

rry)

,slu

rry

fillin

g50

%,

l/d=

2,SS

-cla

d

TN

K-1

,T

NK

-2

flash

tank

K18

[m3]

2.81

120%

13.

3763

63re

side

nce

tim

e:20

min

(liq

uid)

,liq

uid

fillin

g50

%,

l/d=

2,SS

-cla

d

TN

K-1

,T

NK

-2

plug

scre

wfe

eder

K2

[m3/s

]66

6712

0%2

4000

431

861

34ba

rP

SF-1

HT

filte

rpr

ess

K21

[m3/h

]18

.15

120%

121

.78

1324

1324

1)(F

P-1

,FP

-2)

plus

50%

boile

rK

25[k

W]

585

314%

118

3822

922

9w

ood

boile

rw

ith

addi

tion

albu

rner

for

torr

efac

tion

gas

WB

-1,p

lus

20%

for

addi

tion

alga

sbu

rner

drie

rH

XK

32[m

2]

34.9

611

5%1

40.2

193

93SS

/C

SH

X-1

slur

ryH

XK

3[m

2]

8.86

115%

110

.19

181

181

spir

alpl

ate,

SS/

SSH

XS-

1

offga

sH

XK

26[m

2]

2.01

115%

12.

3134

34SS

/C

SH

X-1

air

preh

eate

rK

12[m

2]

2.89

115%

13.

3341

41SS

/C

SH

X-1

coal

cool

erK

37[m

2]

0.25

115%

10.

2813

13ai

r-co

oled

,SS

AC

O-1

cond

ense

rK

28[m

2]

4.37

115%

15.

0269

69ai

r-co

oled

,SS

AC

O-1

cond

ensa

teco

oler

K24

[m2]

11.3

211

5%1

13.0

211

311

3ai

r-co

oled

,SS

AC

O-1

purg

eco

oler

K33

[m2]

19.4

711

5%1

22.3

813

713

7ai

r-co

oled

,SS

AC

O-1

offga

sco

oler

K42

[m2]

0.00

115%

10.

000

0ai

r-co

oled

,SS

AC

O-1

pum

pK

27[k

W]

1.82

220%

22.

0115

30C

SP

-1

rec.

pum

pK

40[k

W]

1.87

110%

12.

0515

15C

SP

-1

drie

rfa

nK

38[m

3/s

]2.

5811

0%1

2.84

33

AF

-1

biom

ass

scre

wco

nvey

or[k

g/h_

ar]

6667

120%

180

0019

19l=

43m

SCC

-1

276


com

pone

ntno

.un

itX

sim

f dn

Xd/

nC

BM

/

[k€

]

CB

M

[k€

]

spec

ifica

tion

sco

stfu

ncti

on

bioc

oalb

elt

drie

rK

22[k

g ev/h

]35

512

0%1

426

473

473

plus

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277


B.4.11 HTC-3.90 with AD of the waste water

B.4.11.1 Simulation data

Table B.34: Flow stream data for the anaerobic digestion in HTC-3.90.

press water condensate substrate, total digestate biogas

temperature [°C] 43.4 37.0 41.1 37.5

CO2 [kg/h] 0.0 1.6 1.6 2.6 59.4

CO [kg/h] 0.0 0.5 0.5 0.1 0.4

H2 [kg/h] 0.0 0.0 0.0 0.0 0.0

H2O [kg/h] 3916.8 2226.2 6143.0 6126.7 4.2

CH4 [kg/h] 0.0 0.0 0.0 0.1 32.6

H2S [kg/h] 0.0 0.7 0.7 0.6 0.1365

C2H4O2 [kg/h] 56.3 8.3 64.6 12.9 0.0

CH2O2 [kg/h] 3.9 0.7 4.6 0.9 0.0

ash [kg/h] 50.2 0.0 50.2 50.2 0.0

TOMres [kg/h] 55.9 0.0 55.9 0.0 0.0

digestate (d.b.) [kg/h] 0.0 0.0 0.0 30.2 0.0

COD [g/l] 53.230 4.089 35.422 14.439

oDM [g/l] 29.539 4.060 20.306 7.173

dry matter content (w%) [–] 4.1% 0.4% 2.8% 1.5%

oDM content (w%, d.b.) [–] 69.8% 100.0% 71.3% 46.7%

B.4.11.2 Cost data

Digester costs are calculated with cost function ADF-1 from Table A.37, for the dimensionssee section B.3.4.

Offsite costs are assumed to be identical to HTC-3.00, additional land cost for the digestersis neglected.

The results of the investment cost estimate are given in Table B.28.

B.4.12 Flowsheet designs of the integrated HTC and CHP plants CHPB-3.1and CHPB-3.3

In the plant design CHPB-1 (Figure B.7), the only connection between the CHP and HTCprocesses is a steam turbine extraction (98) which provides thermal energy for raisingsteam (K54) for the HTC reactor and for an additional slurry preheater (K48). The HTCplant has otherwise the same design as HTC-3.00. The CHP plant consists of a stokerboiler (K56), steam turbine (K49, K50), heating condenser (K53), feedwater pump (K52)and air preheater (K51). 2.7% of the steam are extracted from the turbine at a pressureof 34.7 bar for the HTC plant. The returning condensate (85) is mixed with the steamturbine exhaust steam (100) and led to the heating condenser (K53).

Simulation case CHPB-3 (Figure B.8) is the same configuration as CHPB-2, but employs asuperheated steam drier. Some low pressure steam is re-compressed as fluidization steam

278


for the drier (61). The exhaust steam of the drier is used for district heat production(K57).

biocoal

condensate

reactor

pelletpress

flue gas

filter pressair

76

50

14

47

19 20 21 22 23

6

15

13

96

54

53

1639

37 4

10

7

77

85 100

5

5

52

34

3235

8011 86

12

87

82

81

56

62

92

41

97

61

28

63

64

75

60

29

63

95

95

93

93

9

46

91

55

8340

8

57

44

36

5868+99

30

42

K2 K3 K4

K48

K5

K28

K24

K45

K6 K7 K8 K9 K10

K49K50

K51

K52K53

K55

K54

K11

K16 K15

K29

K19

K31

K21 K37

K46K22

K20

K18

K41

K40

K36

K27

K38K32

K26

K56

K42

K13

K14

K17

17 1871

drier

A

A

B

D

D

B

biomass

27

1 K1

K33

3

2

W1 W2 W4

W9W7

W5

W11

W8

W3

wood

steamturbine

heatingcondenser

exhaustgas

air

ash 72

73

98

78

48

67 51

74

79W14+W15

W16

G

C

C

wastewater

aeration

discharge

94

W18

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

K44

Figure B.7: Flowsheet of CHPB-3.1.

condensate

reactor

filter press

76

4750

14

11

126

96

54

53

4

10

7

77

85

83

83

100

5

5

80 87

569286

41

97

60

29

19

95

95

93

93

9

46

91

55

8

57

44

36

5830

42

K2 K3 K4

K48

K58

K28

K24

K45

K49K50

K52K53

K57

K60

K55

K54K11

K16K19

K31

K21 K37

K20

K36

K27

K26

K41

K56

K42

K13

K14

17 1871

A

A

B

B

biomass

27

1 K1

K33

3

2 W1

W7

W3

wood

steamturbine

hotwater

heatingcondenser

heatingcondenser

ash 72

73

98

78

48

15 45

67 51

20

22

21

23

43

19

74

79W14+W15

W16

G

C

C

air

52

16

16

13

13

81

81

82

82

E

E

F

F

G

G

66

5934

3235 K51

exhaustgas

25

26

61

6864

99K22

K38

SSD

D

D

VOCloss

W8

K61

biocoalpelletpress

75K46

W9

K44wastewater

aeration

discharge

94

W18

biomass

steam

combustiblegas

flue gaselectricity

biocoal

liquid water

air

Figure B.8: Flowsheet of CHPB-3.3.

279


B.4.13 Simulation data from CHPB-3.2

Table B.35: CHPB-3.2 flow stream data.

stream no. 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 21 22T [°C] 15.0 90.0 56.6 98.8 60.0 219.0 232.0 152.0 50.0 213.0 152.0 148.1 140.0 152.0 55.0 15.0 58.3 90.0 121.2 55.0p [bar] 1.0 1.0 1.0 27.2 3.0 25.2 29.0 5.0 1.0 27.2 5.0 5.0 1.0 5.0 5.0 1.0 27.2 27.2 1.0 5.0m [kg/h] 6667 6887 13554 13554 185 15745 3252 13557 2335 16805 1686 1686 27300 2188 413258 27300 13554 13554 77204 16200H [kW] -23536 -29159 -52695 -52097 -453 -60353 -11856 -52348 -10182 -63953 -6171 -7141 246 -8005 -1808000 -724 -52666 -52222 -72965 -70873O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 6319.00 0.00 0.00 6319.00 0.00 0.00 5065.03 0.00CO2 [kg/h] 0.00 0.00 0.00 0.00 171.26 10.02 0.00 0.04 11.50 0.00 7.69 7.69 0.00 9.97 0.00 0.00 0.00 0.00 12472.47 0.00CO [kg/h] 0.00 0.00 0.00 0.00 6.36 1.67 0.00 0.04 2.07 0.00 1.25 1.25 0.00 1.62 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/h] 0.00 0.00 0.00 0.00 0.69 0.02 0.00 0.00 0.02 0.00 0.02 0.02 0.00 0.02 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 4667 6667 11333 11333 6 14168 3252 12001 2313 14585 1671 1671 173 2167 413258 173 11333 11333 11871 16200CH4 [kg/h] 0.00 0.00 0.00 0.00 0.17 0.14 0.00 0.01 0.18 0.00 0.10 0.10 0.00 0.13 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/h] 0 0 0 0 0 0 0 0 0 0 0 0 20808 0 0 20808 0 0 47788 0H2S [kg/h] 0.00 0.00 0.00 0.00 0.18 0.83 0.00 0.18 0.90 0.00 0.51 0.51 0.00 0.66 0.00 0.00 0.00 0.00 0.00 0.00C2H4O2 [kg/h] 0.00 75.52 75.52 75.52 0.00 147.73 0.00 140.53 6.90 75.52 5.54 5.54 0.00 7.19 0.00 0.00 75.52 75.52 0.00 0.00CH2O2 [kg/h] 0.00 5.26 5.26 5.26 0.00 10.38 0.00 9.75 0.61 5.26 0.48 0.48 0.00 0.63 0.00 0.00 5.26 5.26 0.00 0.00SO2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 7.87 0.00biocoal DM [kg/h] 0.00 0.00 0.00 0.00 0.00 1139.82 0.00 1139.82 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00PGW DM [kg/h] 2000 0 2000 2000 0 0 0 0 0 2000 0 0 0 0 0 0 2000 2000 0 0wood DM [kg/h] 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0TOMres [kg/h] 0.00 73.39 73.39 73.39 0.00 140.09 0.00 140.09 0.00 73.39 0.00 0.00 0.00 0.00 0.00 0.00 73.39 73.39 0.00 0.00ash, diss. [kg/h] 0.00 66.02 66.02 66.02 0.00 126.01 0.00 126.01 0.00 66.02 0.00 0.00 0.00 0.00 0.00 0.00 66.02 66.02 0.00 0.00E [MW] 10.127 1.238 11.339 11.451 0.067 11.369 0.989 10.741 0.082 12.268 0.421 0.109 0.199 0.546 7.001 0.034 11.350 11.424 2.470 0.274stream no. 23 27 28 29 30 32 34 35 36 41 42 44 45 46 47 48 50 51 52 53T [°C] 95.0 90.0 18.7 40.0 100.0 190.0 159.0 151.5 40.0 100.0 220.0 50.0 95.0 15.8 152.0 55.0 152.0 65.1 15.0 90.0p [bar] 5.0 1.0 1.1 1.0 1.0 1.0 1.0 1.0 1.0 1.0 25.2 1.0 5.0 33.0 5.0 5.0 5.0 1.4 1.0 1.0m [kg/h] 16200 5473 27800 1282 1460 77204 62691 77204 2335 254 15745 2335 413258 3252 280 429458 221 32250 62691 1925H [kW] -70119 -23174 -708 -1349 -5370 -71245 905 -72211 -10208 -935 -60332 -10182 -1788700 -14365 -1025 -1878800 -810 -140720 -1663 -4083O2 [kg/h] 0.00 0.00 6434.73 0.00 0.00 5065.03 14510.72 5065.03 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 14510.72 0.00CO2 [kg/h] 0.00 0.00 0.00 0.00 10.00 12472.47 0.00 12472.47 11.50 1.74 10.02 11.50 0.00 0.00 1.28 0.00 1.01 0.00 0.00 0.00CO [kg/h] 0.00 0.00 0.00 0.00 1.65 0.00 0.00 0.00 2.07 0.29 1.67 2.07 0.00 0.00 0.21 0.00 0.16 0.00 0.00 0.00H2 [kg/h] 0.00 0.00 0.00 0.00 0.02 0.00 0.00 0.00 0.02 0.00 0.02 0.02 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 16200 5298 176 127 1443 11871 397 11871 2313 251 14168 2313 413258 3252 277 429458 219 32250 397 760CH4 [kg/h] 0.00 0.00 0.00 0.00 0.14 0.00 0.00 0.00 0.18 0.02 0.14 0.18 0.00 0.00 0.02 0.00 0.01 0.00 0.00 0.00N2 [kg/h] 0 0 21189 0 0 47788 47783 47788 0 0 0 0 0 0 0 0 0 0 47783 0H2S [kg/h] 0.00 0.00 0.00 0.00 0.73 0.00 0.00 0.00 0.90 0.13 0.83 0.90 0.00 0.00 0.08 0.00 0.07 0.00 0.00 0.00C2H4O2 [kg/h] 0.00 60.02 0.00 1.44 3.58 0.00 0.00 0.00 6.90 0.62 147.73 6.90 0.00 0.00 0.92 0.00 0.73 0.00 0.00 8.61CH2O2 [kg/h] 0.00 4.18 0.00 0.10 0.35 0.00 0.00 0.00 0.61 0.06 10.38 0.61 0.00 0.00 0.08 0.00 0.06 0.00 0.00 0.60SO2 [kg/h] 0.00 0.00 0.00 0.00 0.00 7.87 0.00 7.87 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00biocoal DM [kg/h] 0.00 0.00 0.00 1139.82 0.00 0.00 0.00 0.00 0.00 0.00 1139.82 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1139.82PGW DM [kg/h] 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0wood DM [kg/h] 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0TOMres [kg/h] 0.00 58.33 0.00 5.86 0.00 0.00 0.00 0.00 0.00 0.00 140.09 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 8.37ash, diss. [kg/h] 0.00 52.47 0.00 7.52 0.00 0.00 0.00 0.00 0.00 0.00 126.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 7.52E [MW] 0.404 0.984 0.058 8.227 0.270 3.032 0.565 2.694 0.080 0.047 11.378 0.082 10.302 0.048 0.070 7.276 0.055 0.594 0.078 8.302stream no. 54 55 56 57 58 59 60 61 62 63 66 67 68 71 72 73 74 75 76 77T [°C] 90.0 135.0 15.0 110.0 149.6 160.0 40.7 100.0 15.0 40.0 15.0 109.9 40.9 70.3 800.0 66.4 95.0 15.0 99.8 160.0p [bar] 1.0 32.7 1.0 5.0 5.0 1.0 1.0 1.0 1.0 1.0 1.0 1.4 1.0 27.2 1.0 94.0 5.0 1.0 27.2 34.7m [kg/h] 12360 3252 3252 221 15745 23391 1282 816 27800 816 23391 32250 28441 13554 234 32250 429458 1282 13554 2963H [kW] -52333 -13925 -14370 -950 -61633 345 -1349 -3003 -738 -3565 -621 -121470 -2907 -52499 28 -140600 -1858800 -1363 -52082 -12595O2 [kg/h] 0.00 0.00 0.00 0.00 0.00 5414.15 0.00 0.00 6434.73 0.00 5414.15 0.00 6434.73 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 0.00 0.00 1.01 10.02 0.00 0.00 5.59 0.00 5.59 0.00 0.00 0.02 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO [kg/h] 0.00 0.00 0.00 0.16 1.67 0.00 0.00 0.92 0.00 0.92 0.00 0.00 0.02 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/h] 0.00 0.00 0.00 0.00 0.02 0.00 0.00 0.01 0.00 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 11965 3252 3252 219 14168 148 127 807 176 807 148 32250 809 11333 0 32250 429458 127 11333 2963CH4 [kg/h] 0.00 0.00 0.00 0.01 0.14 0.00 0.00 0.08 0.00 0.08 0.00 0.00 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/h] 0 0 0 0 0 17829 0 0 21189 0 17829 0 21189 0 0 0 0 0 0 0H2S [kg/h] 0.00 0.00 0.00 0.07 0.83 0.00 0.00 0.41 0.00 0.41 0.00 0.00 0.11 0.00 0.00 0.00 0.00 0.00 0.00 0.00C2H4O2 [kg/h] 135.54 0.00 0.00 0.73 147.73 0.00 1.44 2.00 0.00 2.00 0.00 0.00 7.17 75.52 0.00 0.00 0.00 1.44 75.52 0.00CH2O2 [kg/h] 9.44 0.00 0.00 0.06 10.38 0.00 0.10 0.19 0.00 0.19 0.00 0.00 0.50 5.26 0.00 0.00 0.00 0.10 5.26 0.00SO2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00biocoal DM [kg/h] 0.00 0.00 0.00 0.00 1139.82 0.00 1139.82 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1139.82 0.00 0.00PGW DM [kg/h] 0 0 0 0 0 0 0 0 0 0 0 0 0 2000 0 0 0 0 2000 0wood DM [kg/h] 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0TOMres [kg/h] 131.72 0.00 0.00 0.00 140.08 0.00 5.86 0.00 0.00 0.00 0.00 0.00 0.00 73.39 0.00 0.00 0.00 5.86 73.39 0.00ash, diss. [kg/h] 118.49 0.00 0.00 0.00 126.01 0.00 7.52 0.00 0.00 0.00 0.00 0.00 0.00 66.02 129.90 0.00 0.00 7.52 66.02 0.00E [MW] 2.221 0.121 0.045 0.011 10.875 0.213 8.227 0.151 0.035 0.028 0.029 5.283 0.091 11.374 1.020 0.684 10.705 8.226 11.454 0.140stream no. 78 79 80 81 82 83 85 87 91 92 93 94 95 96 97 98 99 100T [°C] 500.0 15.0 218.3 15.0 200.0 129.5 80.7 69.7 90.0 100.0 36.8 40.0 60.0 100.0 68.3 386.2 40.7 109.9p [bar] 80.0 1.0 23.2 1.0 1.0 5.0 34.7 23.2 32.7 1.0 1.0 1.0 3.0 1.0 1.0 34.7 1.0 1.4m [kg/h] 32250 14562 1060 12000 12000 280 2963 1060 3252 389 389 5473 876 14285 254 2963 3 29287H [kW] -112710 -42598 -3629 -318 315 -1195 -12872 -4262 -14096 -1432 -1701 -23477 -3814 -56264 -1102 -10527 -2 -108600O2 [kg/h] 0.00 0.00 0.00 2777.58 2777.58 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/h] 0.00 0.00 172.76 0.00 0.00 1.28 0.00 172.76 0.00 2.67 2.67 0.00 1.50 0.02 1.74 0.00 0.00 0.00CO [kg/h] 0.00 0.00 6.78 0.00 0.00 0.21 0.00 6.78 0.00 0.44 0.44 0.00 0.42 0.02 0.29 0.00 0.00 0.00H2 [kg/h] 0.00 0.00 0.69 0.00 0.00 0.00 0.00 0.69 0.00 0.01 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/h] 32250 7281 876 76 76 277 2963 876 3252 385 385 5298 870 12725 251 2963 0 29287CH4 [kg/h] 0.00 0.00 0.22 0.00 0.00 0.02 0.00 0.22 0.00 0.04 0.04 0.00 0.05 0.01 0.02 0.00 0.00 0.00N2 [kg/h] 0 0 0 9146 9146 0 0 0 0 0 0 0 0 0 0 0 0 0H2S [kg/h] 0.00 0.00 0.36 0.00 0.00 0.08 0.00 0.36 0.00 0.19 0.19 0.00 0.17 0.11 0.13 0.00 0.00 0.00C2H4O2 [kg/h] 0.00 0.00 3.33 0.00 0.00 0.92 0.00 3.33 0.00 0.95 0.95 60.02 3.32 144.14 0.62 0.00 0.00 0.00CH2O2 [kg/h] 0.00 0.00 0.27 0.00 0.00 0.08 0.00 0.27 0.00 0.09 0.09 4.18 0.27 10.03 0.06 0.00 0.00 0.00SO2 [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00biocoal DM [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1139.82 0.00 0.00 0.00 0.00PGW DM [kg/h] 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0wood DM [kg/h] 0 7281 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0TOMres [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 58.33 0.00 140.08 0.00 0.00 2.51 0.00ash, diss. [kg/h] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 52.47 0.00 126.01 0.00 0.00 0.00 0.00E [MW] 13.549 41.635 0.342 0.015 0.159 0.015 0.066 0.107 0.078 0.072 0.013 0.940 0.032 10.557 0.010 1.056 0.026 5.222

[kWel]W1 38.1 slurry pump (K2)W3 5.9 pump (K27)W7 14.5 filter press (K21)W8 34.6 drier fan (K38)W9 64.1 pellet press (K46)W13 13.3 cooler fansW14 -1770.7 HP steam turbine (K49)W15 -4325.6 LP steam turbine (K50)W16 131.9 feedwater pump (K52)W17 5541.1 overall net electricityW18 252.9 waste water aeration

280


B.4.14 Cost data from CHPB-3.1, CHPB-3.2 and CHPB-3.3

Table B.36: Investment cost summary of the integrated HTC and CHP systems.


HTC

HTC reactor [M€] 1.77 1.77 1.76 1.76

slurry pumps [M€] 0.46 0.47 0.29 0.29

flash tanks [M€] 0.53 0.54 0.17 0.17

filter press [M€] 0.74 0.73 0.79 0.79

biocoal drier (incl. HX) [M€] 0.78 0.78 0.78 1.04

pellet press [M€] 0.34 0.34 0.34 0.34

boiler [M€] 0.23 0.00 0.00 0.00

heat exchangers [M€] 0.51 0.56 0.56 0.55

coolers [M€] 0.35 0.34 0.33 0.36

other [M€] 0.80 0.79 0.84 0.86

CBM HTC [M€] 6.51 6.31 5.86 6.16

CHP

boiler [M€] 7.95 8.09 8.09 7.91

steam turbine [M€] 1.89 1.89 1.87 1.83

heating condenser [M€] 0.19 0.18 0.24 0.31

air preheater [M€] 0.37 0.38 0.35 0.38

feedwater pump [M€] 0.24 0.24 0.25 0.24

CBM CHP [M€] 10.63 10.79 10.80 10.68

offsite cost [M€] 2.66 1.96 1.96 1.96

fees & contingencies [M€] 2.57 2.57 2.50 2.53

start-up [M€] 0.73 0.67 0.66 0.66

working capital [M€] 1.70 1.57 1.57 1.54

AFUDC [M€] 2.24 2.16 2.11 2.13

residual value (NPV) [M€] -0.28 -0.22 -0.23 -0.22

TCI [M€] 26.76 25.82 25.23 25.44

281


Table B.37: Annual cost flows for the standalone CHP plant and the integrated plants.

CHP

standalone

HTC

standalone

CHPB-1 CHPB-2 CHPB-3

expenses (A)

carrying charges [k€/a] 2356 1410 3767 3634 3551

labour [k€/a] 538 1021 1559 1254 1254

electricity purchase [k€/a] 299

O&M, material [k€/a] 212 254 466 465 447

other operating cost [k€/a] 70 4 74 76 76

waste incl. transport [k€/a] -689 -689 -689 -689

wood incl. transport [k€/a] 5097 5097 5221 5217

revenues (B)

electricity feed-in -3397 -3397 -3188 -3107

district heat -4879 -4876 -4876 -4877

TRR biocoal (C=A+B) 0 2299 2001 1896 1871

biocoal production cost [€/GJHHV] 12.66 9.48 9.38 9.31

282


Tabl

eB

.38:

Equ

ipm

ent

list

wit

hin

vest

men

tco

sts

ofth

est

anda

lone

CH

Ppl

ant.

com

pone

ntun

itX

sim

f dn

Xd/

nC

BM

/n [k€

]

CB

M

[k€

]

spec

ifica

tion

s,co

mm

ents

cost

func

tion

CH

Pbo

iler

[MW

HH

V]

38.7

120%

146

.479

4879

48p d

=1.

1p s

im=

88ba

r,20

5°C

supe

rhea

ting

WB

-2

stea

mtu

rbin

e[M

Wel

]6.

1911

0%1

6.80

1885

1885

STB

-1,

STB

-2

heat

ing

cond

ense

r[m

2]

343

115%

139

518

718

7C

S/C

uH

X-1

air

preh

eate

r[m

2]

1520

115%

117

4836

936

9ro

tary

air

preh

eate

rA

PH

-1

feed

wat

erpu

mp

[kW

el]

125

220%

213

812

024

0C

SP

-1

tota

lC

BM

1062

9

283


B.4.15 HTC followed by the subsequent combustion in a CHP plant

wetbiomass

conden-sate

exhaustgas

dried biomass

airash

5

3114

16

15

44

40

41

43

42

28

25

3132

24

2 39

13

17

3538 37

3627

26

9

6

10

W1

W5

W2

W3

G

SSD

Figure B.9: Flowsheet of SC-3.1.4.

284


Tabl

eB

.39:

Ene

rgy

bala

nce

(HH

V)

ofH

TC

follo

wed

byth

esu

bseq

uent

com

bust

ion

ina

CH

Ppl

ant

and

dire

ctco

mbu

stio

nof

raw

biom

ass,

SC-1

.0.1

toSC

-3.1

.4.

case

SC-

1.0.

11.

0.2

1.0.

31.

1.1

1.1.

23.

0.1

3.0.

23.

0.3

3.1.

13.

1.2

3.1.

4

HT

Cp

lant

inpu

tsra

wbi

omas

s[M

W]

55.4

0—

—58

.14

—58

.91

——

61.6

8—

—

elec

tric

ity

[MW

]1.

82—

—1.

91—

2.56

——

2.68

——

outp

utbi

ocoa

l[M

W]

44.6

8—

—46

.89

—44

.61

——

46.7

1—

—

CH

Pp

lant

inpu

tsbi

omas

s[M

W]

—54

.76

60.1

2—

54.3

0—

108.

1315

5.47

—10

5.50

48.8

8

bioc

oal

[MW

]44

.68

——

46.8

9—

44.6

1—

—46

.71

——

outp

utel

ectr

icit

y[M

W]

5.92

5.92

8.36

9.67

9.34

5.92

13.5

824

.06

9.67

15.8

07.

78

ther

mal

ener

gy[M

W]

31.9

831

.98

31.9

831

.90

31.9

031

.98

31.9

831

.98

31.9

031

.90

31.9

0

over

all

syst

em

inpu

tsra

wbi

omas

s[M

W]

55.4

054

.76

60.1

258

.14

54.3

058

.91

108.

1315

5.47

61.6

810

5.50

48.8

8

outp

utel

ectr

icit

y(n

et)

[MW

]5.

925.

928.

369.

679.

345.

9213

.58

24.0

69.

6715

.80

7.78

ther

mal

ener

gy[M

W]

31.9

831

.98

31.9

831

.90

31.9

031

.98

31.9

831

.98

31.9

031

.90

31.9

0

CH

Pp

lant

effici

enci

es

elec

tric

al,H

HV

[–]

13.2

%10

.8%

13.9

%20

.6%

17.2

%13

.3%

12.6

%15

.5%

20.7

%15

.0%

15.9

%

ener

geti

c,H

HV

[–]

84.8

%69

.2%

67.1

%88

.6%

76.0

%85

.0%

42.1

%36

.0%

89.0

%45

.2%

81.2

%

elec

tric

al,L

HV

[–]

14.0

%13

.4%

17.3

%21

.8%

21.4

%14

.0%

21.2

%26

.2%

21.8

%25

.3%

26.9

%

ener

geti

c,L

HV

[–]

89.7

%86

.1%

83.5

%93

.8%

94.5

%89

.6%

71.2

%60

.9%

93.9

%76

.4%

137.

2%

285


TableB

.40:E

xergybalance

ofHT

Cfollow

edby

thesubsequent

combustion

ina

CH

Pplant

anddirect

combustion

ofrawbiom

ass,SC

-1.0.1to

SC-3.1.4.

caseSC

-1.0.1

1.0.21.0.3

1.1.11.1.2

3.0.13.0.2

3.0.33.1.1

3.1.23.1.4

inp

uts

biomass,dry

[MW

]58.20

57.5263.16

61.0857.04

62.35114.43

164.5365.28

111.6451.73

moisture

[MW

]0.14

0.140.15

0.150.14

0.400.74

1.060.42

0.720.33

air,HT

C[M

W]

0.18—

—0.19

—0.21

——

0.22—

—

air,CH

P[M

W]

0.080.09

0.660.08

0.090.08

1.944.41

0.081.89

0.08

outp

ut

electricity(net)

[MW

]4.09

5.928.36

7.759.34

3.3613.58

24.066.99

15.807.78

thermal

[MW

]11.70

11.7011.70

5.485.48

11.7011.70

11.705.48

5.485.48

ED

+E

L

HT

Cplant

ED

+E

L[M

W]

14.05—

—14.74

—19.10

——

20.00—

—

boilerE

D[M

W]

25.2434.80

35.4927.24

35.0325.22

69.6193.94

27.1269.53

29.43

drierE

D[M

W]

——

2.72—

——

8.4519.73

—8.24

2.34

steamturbine

ED

[MW

]1.06

1.061.64

1.901.83

1.062.88

5.341.90

3.401.55

EP

Hboiler

exhaust[M

W]

0.712.03

0.610.23

1.140.70

2.831.03

0.222.24

0.44

EP

Hdrier

exhaust[M

W]

——

0.30—

——

0.932.30

—0.90

0.16

CH

P,otherE

D+

EL

[MW

]1.74

2.253.15

4.144.44

1.887.14

11.904.28

8.654.97

totalE

D+

EL

[MW

]42.81

40.1443.92

48.2642.44

47.9791.83

134.2353.52

92.9738.88

exergeticeffi

ciencies

ofselected

comp

onents

boiler[–]

42.6%35.0%

40.7%41.5%

34.8%42.6%

34.2%39.9%

41.7%33.5%

40.3%

CH

Pplant

[–]38.0%

30.5%31.4%

31.1%25.9%

37.9%21.6%

21.0%31.1%

18.6%25.4%

286

B.5 GHG emissions from biofuel production (WP, TOR, HTC, ADM, ADP)

B.5 GHG emissions from biofuel production (WP, TOR, HTC,ADM, ADP)

Table B.41: GHG emissions for biofuel production for WP, TOR, HTC, ADM and ADP,in [kgCO2,eq/GJHHV,biofuel]. The system boundary is the plant gate of the upgradingplant.

feedstockroad

transport

biofuel,road

transport

biofuel,ship

transport

auxiliaryenergy

methaneemis-sions

avoidedGHG,

dumping

total

WP1.0-s-SR 0.467 3.832 4.300WP1.0-m-SR 0.786 3.832 4.618WP1.0-l-SR 1.655 3.832 5.487WP1.0-s-FR 0.357 3.863 3.549 3.832 11.601WP1.0-m-FR 0.692 3.863 3.549 3.832 11.937WP1.0-l-FR 1.399 3.863 3.549 3.832 12.643WP2.0-s-SR 0.384 7.787 8.171WP2.0-m-SR 0.641 7.787 8.427WP3.0-m-SR 0.957 -4.684 -3.726WP3.0-l-SR 2.023 -4.684 -2.660WP3.1-m-SR 0.890 -4.684 -3.794WP3.1-l-SR 1.879 -4.684 -2.805WP3.0-m-FR 0.840 3.863 3.549 -4.684 3.568WP3.0-l-FR 1.704 3.863 3.549 -4.684 4.432WP3.1-m-FR 0.782 3.863 3.549 -4.684 3.510WP3.1-l-FR 1.584 3.863 3.549 -4.684 4.312TOR-1.0-s-SR 0.486 2.523 3.009TOR-1.0-m-SR 0.816 2.523 3.340TOR-1.0-l-SR 1.719 2.523 4.242TOR-1.0-s-FR 0.371 3.296 3.028 2.523 9.218TOR-1.0-m-FR 0.719 3.296 3.028 2.523 9.566TOR-1.0-l-FR 1.453 3.296 3.028 2.523 10.300HTC-1.00-s 0.453 7.300 7.753HTC-1.00-m 0.754 7.300 8.055HTC-2.00-s 1.273 8.567 9.840HTC-2.00-m 4.739 8.567 13.305HTC-3.00-s 3.728 10.256 13.984HTC-3.00-m 10.913 10.256 21.169HTC-3.90-s 3.490 6.609 10.099HTC-4.00-s 2.176 12.043 14.219HTC-4.00-m 9.355 12.043 21.398HTC-5.00-s 0.000 1.344 5.646 20.712 -200.619 -172.916HTC-5.00-m 0.000 1.344 5.646 20.712 -200.619 -172.916ADP-3.0-s 1.145 -15.378 1.586 -12.647ADM-3.0s 1.743 17.013 4.502 23.258ADM-3.1-s 0.940 14.390 4.502 19.832

287


B.6 Combustion of upgraded biofuels in existing power stations

Table B.42: Water content, density, calorific value, energy density and costs for 100 kmroad transport for different types of raw biomass and upgraded biofuels.

water

content

bulk

density

(w.b.)

HHV

(w.b.)

LHV

(w.b.)

energy

density

transport

cost

transport

cost

100 km

[–] [kg/m3] [MJ/kg] [MJ/kg] [GJLHV/m3][€/t/km] [€/GJLHV]

wood chips 50% 290 9.8 7.9 2.28 0.101 1.29

PGW-70, grass 70% 167 5.1 3.0 0.51 0.130 4.28

PGW-50 70% 750 4.3 2.2 1.66 0.166 7.49

MOW 50% 100 8.6 6.7 0.67 0.217 3.25

wood pellets 10% 650 17.6 16.1 10.46 0.080 0.50

biocoal, pellets (wood) 10% 750 24.1 22.9 17.15 0.080 0.35

biocoal, pulverized (wood) 10% 450 24.1 22.9 10.29 0.209 0.91

biocoal, pellets (PGW-70) 10% 750 22.1 21.0 15.76 0.080 0.38

biocoal, pulverized (PGW-70) 10% 450 22.1 21.0 9.46 0.209 0.99

biocoal, pellets (MOW) 10% 750 16.6 15.6 11.72 0.080 0.51

ADP-pellets (grass) 10% 650 15.7 14.2 9.26 0.080 0.56

288

B.6 Combustion of upgraded biofuels in existing power stations

Tabl

eB

.43:

Key

resu

lts

from

sele

cted

biom

ass

upgr

adin

gsc

enar

ios.

scen

ario

feed

stoc

kbi

ofue

ldis

plac

esca

paci

ty

(fee

dsto

ck)

capa

city

(fee

dsto

ck)

biof

uelc

ost

atpo

wer

plan

tga

te

GH

G

redu

ctio

npe

r

unit

ofbi

ofue

l

GH

G

redu

ctio

npe

r

unit

of

feed

stoc

k

GH

G

mit

igat

ion

cost

[MW

HH

V,f

eed

][k

t fee

d,F

M/a

][€

/GJ H

HV

][k

g/G

J HH

V,b

iofu

el][k

g/G

J HH

V,f

eed

][€

/tC

O2,

eq]

HT

C-3

.00-

sP

GW

-70

bitu

min

ous

coal

9.5

46.7

13.2

286

.665

.612

0.32

HT

C-3

.00-

s-G

gras

bitu

min

ous

coal

9.5

46.7

26.1

089

.267

.626

1.05

HT

C-3

.00-

mP

GW

-70

bitu

min

ous

coal

47.5

233.

68.

8979

.460

.176

.63

HT

C-3

.90-

sP

GW

-70

bitu

min

ous

coal

9.5

46.7

12.4

490

.573

.210

6.53

HT

C-2

.00-

sP

GW

-50

bitu

min

ous

coal

9.5

28.0

14.3

490

.770

.612

7.20

HT

C-2

.00-

mP

GW

-50

bitu

min

ous

coal

47.5

140.

29.

7587

.267

.979

.67

HT

C-4

.00-

sM

OW

bitu

min

ous

coal

7.9

46.7

6.20

84.1

60.2

40.9

1

HT

C-4

.00-

mM

OW

bitu

min

ous

coal

39.4

233.

60.

8776

.955

.1-2

4.56

HT

C-5

.00-

sE

FB

bitu

min

ous

coal

11.9

40.0

12.0

227

2.6

210.

033

.92

HT

C-5

.00-

m*

EF

Bbi

tum

inou

sco

al26

.690

.010

.52

272.

621

0.0

28.4

2

HT

C-1

.00-

sw

ood,

SRbi

tum

inou

sco

al10

.928

.019

.88

92.6

74.7

184.

44

HT

C-1

.00-

mw

ood,

SRbi

tum

inou

sco

al54

.314

0.2

14.7

092

.374

.512

8.93

HT

C-3

.00-

m-P

LP

GW

-70

pulv

eriz

edlig

nite

47.5

233.

69.

3385

.064

.446

.57

WP

-1.0

-m-S

Rw

ood,

SRbi

tum

inou

sco

al64

.816

7.3

9.53

91.5

76.7

74.6

7

WP

.1.0

-l-S

Rw

ood,

SRbi

tum

inou

sco

al32

4.0

836.

39.

0390

.676

.069

.86

WP

-1.0

-m-F

Rw

ood,

FR

bitu

min

ous

coal

64.8

167.

37.

6384

.270

.658

.64

WP

-1.0

-l-F

Rw

ood,

FR

bitu

min

ous

coal

324.

083

6.3

7.12

83.5

70.0

53.0

2

TO

R-1

.0-m

-SR

woo

d,SR

bitu

min

ous

coal

64.2

165.

710

.12

95.1

76.3

77.5

0

TO

R-1

.0-l

-SR

woo

d,SR

bitu

min

ous

coal

320.

982

8.4

9.48

94.2

75.6

71.4

6

TO

R-1

.0-m

-FR

woo

d,F

Rbi

tum

inou

sco

al64

.216

5.7

7.70

88.8

71.3

55.6

7

TO

R-1

.0-l

-FR

woo

d,F

Rbi

tum

inou

sco

al32

0.9

828.

47.

0588

.170

.748

.77

AD

P-3

.0gr

asbi

tum

inou

sco

al9.

546

.722

.97

107.

770

.116

9.52

AD

M-3

.0gr

asna

tura

lgas

9.5

46.7

37.8

634

.214

.691

7.38

AD

M-3

.1gr

asna

tura

lgas

9.5

46.7

21.4

937

.729

.839

7.61

289

C Data of the BECCS plants

C.1 Syngas production

289

287

201

304

308

234305

293 299

306, 236

276 277

238 233292

208

288 290

211 213

301239

300

254280

257259

260253

214 218 221 224223 226

271

227

261 265262266264

269263

270

295

272273 281

222 225

230

231294

220

219

255, 256

215

216 217

207

274275

252

244

247

250

drier

feedingsystem

coldbox

exhaustgas

wood

air

air N2

N2

O2

CO2

N2

gasifier+particlesremoval

bed 1

bed 2

sulfur

AGR

condensate

water

water

ATR shift shift

W1

W3

W2

W12

W9

W5

W4

solid fuel

steam

syngas

flue gas

electricity

CO2

liquid water

air, O , N2 2

K5

K6

K23

K32

K27 K28

K22 K21K1

K33

K7 K9 K12 K15K13

K16

K24

K17

K10

K29 K8

K11 K14

K19

K20

235condensate

205 298

209ashsand

dolomite

W10K30

cleangas

LP

MP

MP

HP HP LP

LP

LP

LP

HPHP

Figure C.1: Flowsheet of the syngas production process with autothermal reforming FB-wood-3.

291

Chapter C Data of the BECCS plants

TableC.1:

Energy

balancefor

thesyngas

production,normalized

to100

MW

HH

Vgasifier

fuel.N

egativevalues

indicateoutputs.

EF

EF

EF

EF

FB

FB

FB

FB

HT

C-1

HT

C-102

TO

R-1

wood-1

wood-1

wood-2

wood-3

WP

-1

pretreatm

entp

rocess(in

pu

ts)

biomass

[MW

HH

V]

123.98129.46

124.56100.00

100.00100.00

100.00119.34

electricity[M

Wel ]

4.084.54

1.410.00

0.000.00

0.002.14

syngas

prod

uction

process

gasifierfeed

[MW

HH

V]

100.00100.00

100.00100.00

100.00100.00

100.00100.00

cleangas,net

[MW

HH

V]

-77.05-79.01

-72.41-68.21

-64.89-76.58

-77.51-64.56

electricity(net

total)[M

Wel ]

9.939.70

10.389.48

5.626.85

7.734.61

ASU

&O

2com

pressor[M

Wel ]

4.074.04

4.054.25

1.742.88

2.741.78

AG

R[M

Wel ]

1.551.57

1.511.40

0.860.87

0.900.86

CO

2com

pressing[M

Wel ]

1.681.68

1.661.67

1.992.01

2.071.99

milling

&pressurizing

[MW

el ]2.42

2.232.94

0.920.05

0.050.05

0.05

drierfan

[MW

el ]—

——

1.021.05

1.051.05

—

compressors

[MW

el ]0.21

0.200.22

0.230.38

0.060.92

0.38

SGexpander

[MW

el ]—

——

—-0.46

-0.08—

-0.46

thermalenergy

(nettotal)

[MW

th ]-10.54

-8.02-16.77

-3.77-4.62

7.107.58

-24.47

netsteam

/hotw

ater[M

Wth ]

-10.54-8.02

-16.77-3.77

11.169.78

7.58-8.73

reformer

exhaustgas

[MW

th ]—

——

—-15.78

-2.68—

-15.74

292


Table C.2: Exergy balance for the syngas production, normalized to 100 MWex of gasifierfeed exergy.


HTC-1 HTC-1.02 TOR-1 wood-1 wood-1 wood-2 wood-3 WP-1

pretreatment

raw biomass [MWex] 125.67 131.31 124.88 — — — — 120.18

electricity [MWex] 3.94 4.39 1.35 — — — — 2.05

air and water [MWex] 0.69 0.67 0.34 — — — — 0.33

syngas production

fuelsgasification feed [MWex] 100.00 100.00 100.00 100.00 100.00 100.00 100.00 100.00

electricity [MWex] 9.58 9.37 9.91 9.06 5.37 6.55 7.39 4.42

gasification agents

air

[MWex] 6.41 7.06 4.82 4.61 6.04 5.99 5.11 4.85

productssyngas [MWex] 64.15 65.82 59.64 56.28 53.67 63.30 63.80 53.51

CO2 [MWex] 6.04 6.04 5.92 5.97 5.80 5.87 6.03 5.82

steam [MWex] 12.10 11.82 13.01 9.25 4.45 4.44 4.84 8.80

reformer exhaust gas [MWex] — — — — 6.92 1.18 — 6.91

exergy lossesslag / ash [MWex] 0.39 0.39 0.45 0.44 2.10 2.10 2.10 2.11

drier exhaust [MWex] 0.09 0.09 — 2.04 2.05 2.05 2.05 —

other [MWex] 1.42 1.37 1.62 0.59 0.25 0.35 0.33 0.26

exergy destructiongasifier [MWex] 16.72 15.88 19.08 20.08 16.23 18.47 16.23 16.55

reformer [MWex] — — — — 7.74 1.28 4.13 7.74

gas quench & rec.

compressor

[MWex] 1.21 1.14 1.43 1.57 — — — —

Shift [MWex] 2.75 2.98 2.22 1.79 0.93 1.58 0.81 0.92

ASU [MWex] 2.73 2.71 2.69 2.82 1.20 1.99 2.05 1.23

wood drier & mill [MWex] 0.76 0.70 0.76 4.73 4.67 4.67 4.67 —

steam and HW

generation

[MWex] 3.54 3.42 3.83 4.02 2.72 2.74 2.87 2.72

AGR & CO2-comp. [MWex] 3.16 3.18 3.06 2.95 2.26 2.29 2.38 2.27

other [MWex] 0.93 0.90 1.02 1.14 0.47 0.29 0.23 0.46

exergetic efficiency

syngas production [–] 71.0% 71.9% 68.5% 62.9% 63.6% 66.4% 66.4% 68.7%

conversion chain [–] 56.3% 54.8% 55.6% 62.9% 63.6% 66.4% 66.4% 56.9%

293


Table C.3: EF-HTC-1 flow stream data.

stream no. 1 2 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19T [°C] 15.0 71.9 1550.0 900.3 395.2 310.3 310.3 323.6 310.3 164.1 275.0 254.2 418.0 360.9 345.1 246.6 342.2 177.6p [bar] 1.0 1.0 39.0 39.0 37.7 37.5 37.5 40.0 37.5 36.0 35.7 35.7 33.7 33.6 33.5 33.3 31.3 30.8m [kg/s] 93.1 88.4 169.7 378.6 378.6 378.6 209.0 209.0 169.7 187.7 187.7 303.0 303.0 303.0 303.0 303.0 303.0 303.0H [MW] -381.0 -298.1 -414.8 -1352.5 -1658.0 -1706.8 -941.9 -937.7 -764.9 -1040.5 -1008.0 -2518.2 -2521.5 -2553.9 -2562.8 -2618.1 -2618.1 -2711.4O2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 0.00 0.00 19.40 43.29 43.29 43.29 23.89 23.89 19.40 19.39 19.39 19.39 129.20 129.20 129.20 129.20 191.42 191.42CO [kg/s] 0.00 0.00 116.48 259.91 259.91 259.91 143.44 143.44 116.48 116.48 116.48 116.48 46.59 46.59 46.59 46.59 6.99 6.99H2 [kg/s] 0.00 0.00 4.08 9.09 9.09 9.09 5.02 5.02 4.08 4.08 4.08 4.08 9.11 9.11 9.11 9.11 11.96 11.96H2O [kg/s] 9.04 4.42 15.02 33.52 33.52 33.52 18.50 18.50 15.02 33.06 33.06 148.33 103.38 103.38 103.38 103.38 77.91 77.91CH4 [kg/s] 0.00 0.00 0.00 0.01 0.01 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.00 0.00 14.65 32.70 32.70 32.70 18.04 18.04 14.65 14.65 14.65 14.65 14.65 14.65 14.65 14.65 14.65 14.65H2S [kg/s] 0.00 0.00 0.05 0.11 0.11 0.11 0.06 0.06 0.05 0.05 0.05 0.05 0.05 0.05 0.05 0.05 0.05 0.05SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 84.01 84.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.0 0.7 354.3 446.9 237.6 211.1 116.5 120.2 94.6 92.9 106.1 221.2 245.3 226.9 222.0 194.5 208.1 165.0ECH [MW] 2332.0 2331.8 1614.3 3602.3 3602.3 3602.3 1988.0 1988.0 1614.3 1615.2 1615.2 1620.9 1561.3 1561.3 1561.3 1561.3 1535.5 1535.5

stream no. 20 21 26 27 28 30 31 32 33 34 35 36 37 39 43 44 45 46T [°C] 157.4 157.4 38.0 219.0 253.2 25.8 150.2 241.1 71.9 165.0 40.8 254.7 165.2 1200.0 334.3 335.2 160.0 256.8p [bar] 30.7 30.7 30.5 43.0 43.0 1.1 1.0 1.0 1.0 7.0 7.0 43.0 43.0 39.0 144.8 137.3 47.8 44.5m [kg/s] 303.0 275.5 225.3 65.5 5.3 32.0 32.0 32.0 36.6 1.6 1.6 1.5 1.5 1.5 279.4 279.4 23.0 23.0H [MW] -2781.4 -2358.8 -1741.3 11.9 -69.6 -0.3 3.9 6.9 -60.4 -21.1 -25.3 -19.2 -22.3 4.9 -4031.6 -3726.0 -352.1 -303.3O2 [kg/s] 0.00 0.00 0.00 62.61 0.00 0.77 0.77 0.77 0.77 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 191.42 191.34 191.25 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO [kg/s] 6.99 6.99 6.99 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 11.96 11.95 11.95 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 77.91 50.50 0.42 0.00 5.32 0.02 0.02 0.02 4.64 1.60 1.60 1.46 1.46 0.00 279.39 279.39 23.01 23.01CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 14.65 14.65 14.65 2.88 0.00 31.21 31.21 31.21 31.21 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/s] 0.05 0.05 0.05 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1.37 0.00 0.00 0.00 0.00EPH [MW] 141.0 137.3 90.1 21.6 6.1 0.1 0.9 2.0 1.4 1.3 0.0 1.5 0.2 3.3 144.8 305.4 2.8 24.5ECH [MW] 1535.5 1534.1 1531.9 6.9 0.3 0.6 0.6 0.6 0.8 0.1 0.1 0.1 0.1 5.7 14.0 14.0 1.1 1.1


stream no. 71 72 73 74 78 79 83 84 85 87 88 89 94 95 97 98 102 103T [°C] 253.2 247.3 160.0 256.8 15.0 30.0 20.0 20.0 20.0 20.0 30.0 30.0 20.0 20.0 34.0 49.0 34.0 49.0p [bar] 43.0 31.1 47.8 44.5 1.0 4.6 1.0 1.0 1.0 1.0 60.0 60.0 1.0 1.0 27.2 1.5 27.2 23.3m [kg/s] 115.3 303.0 25.5 25.5 294.1 294.1 65.5 65.5 0.0 69.3 69.3 57.3 32.0 125.6 52.2 0.3 0.0 51.8H [MW] -1510.2 -2672.2 -390.2 -336.1 -28.1 -25.2 -0.3 -0.3 0.0 -1.0 -1.0 -0.9 -0.4 -1.8 -197.3 -1.9 0.0 -462.6O2 [kg/s] 0.00 0.00 0.00 0.00 68.08 68.08 62.64 62.61 0.03 1.66 1.66 1.38 0.77 3.02 0.00 0.00 0.00 0.00CO2 [kg/s] 0.00 191.42 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 18.93 0.20 0.00 51.64CO [kg/s] 0.00 6.99 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 6.80 0.00 0.00 0.06H2 [kg/s] 0.00 11.96 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 11.74 0.00 0.00 0.06H2O [kg/s] 115.27 77.91 25.50 25.50 1.86 1.86 0.00 0.00 0.00 0.04 0.04 0.04 0.02 0.08 0.22 0.01 0.00 0.06CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.00 14.65 0.00 0.00 224.20 224.20 2.89 2.88 0.00 67.57 67.57 55.90 31.21 122.53 14.55 0.00 0.00 0.03H2S [kg/s] 0.00 0.05 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.05 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 131.5 181.3 3.1 27.1 0.0 37.1 0.0 0.0 0.0 0.0 24.2 20.0 0.0 0.0 55.8 0.0 0.0 8.8ECH [MW] 5.8 1535.5 1.3 1.3 1.3 1.3 6.9 6.9 0.0 1.3 1.3 1.1 0.6 2.4 1440.3 1.2 0.1 30.9

stream no. 104 105 108 109 201 202 203 204 205 206 207 208T [°C] 49.0 49.0 40.0 40.0 166.1 100.0 166.1 100.0 160.0 166.8 160.0 256.8 [MWel]p [bar] 9.2 2.9 110.0 50.6 7.2 7.2 7.2 7.2 7.4 7.2 47.8 44.5 W1 4.43 recycle compressor (K3)m [kg/s] 43.2 77.8 172.8 0.0 4.6 4.6 21.7 21.7 0.0 0.0 0.0 0.0 W3 16.75 coal millH [MW] -385.0 -692.6 -1571.9 -0.4 -60.2 -70.9 -286.5 -337.3 -0.3 -0.2 -0.2 -0.2 W4 6.55 ASU cold box (K26)O2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 W5 51.12 ASU air compressor (K27)CO2 [kg/s] 43.03 77.45 172.12 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 W6 33.98 O2 compressor (K28)CO [kg/s] 0.05 0.08 0.19 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 W7 0.20 N2 compressor (K30)H2 [kg/s] 0.05 0.10 0.21 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 W8 37.71 N2 compressor (K29)H2O [kg/s] 0.05 0.09 0.17 0.02 4.55 4.55 21.68 21.68 0.02 0.02 0.01 0.01 W10 34.87 AGR (K25)CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 W11 37.79 CO2 compressor (K31)N2 [kg/s] 0.02 0.04 0.10 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 W12 223.40 total electricity consumptionH2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 5.3 4.7 37.8 0.0 3.8 0.2 18.1 1.0 0.0 0.0 0.0 0.0ECH [MW] 25.8 46.4 103.1 0.0 0.2 0.2 1.1 1.1 0.0 0.0 0.0 0.0

294


Table C.4: FB-wood-1 flow stream data.

stream no. 201 205 207 208 209 211 213 214 215 216 217 218 219 220 221 222 223 224 225 226T [°C] 15.0 899.8 359.6 163.3 15.0 900.0 950.0 480.0 480.0 400.0 400.0 350.0 348.3 350.0 475.1 328.1 200.0 220.8 170.0 129.6p [bar] 1.0 1.0 37.3 37.3 33.3 31.6 29.6 28.7 28.7 28.5 26.5 26.4 26.4 26.9 24.1 24.1 23.6 21.3 21.0 21.0m [kg/s] 47.8 2.0 9.8 5.7 0.7 41.5 41.5 41.5 41.5 41.5 41.5 41.5 50.0 8.5 50.0 50.0 50.0 50.0 50.0 50.0H [MW] -507.1 -16.4 -125.6 0.7 -8.6 -305.7 -240.4 -281.5 -281.5 -288.1 -288.1 -292.2 -401.3 -109.1 -401.6 -416.8 -429.6 -429.6 -434.6 -441.5O2 [kg/s] 0.00 0.00 0.00 5.45 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 20.20 20.20 20.20 20.20 20.20 20.20 20.20 20.20 0.00 35.71 35.71 35.71 38.04 38.04 38.04CO [kg/s] 0.00 0.00 0.00 0.00 0.00 4.06 11.62 11.62 11.62 11.62 11.62 11.62 11.62 0.00 1.74 1.74 1.74 0.26 0.26 0.26H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.38 1.99 1.99 1.99 1.99 1.99 1.99 1.99 0.00 2.70 2.70 2.70 2.81 2.81 2.81H2O [kg/s] 23.88 0.00 9.78 0.00 0.00 11.30 6.45 6.45 6.45 6.45 6.45 6.45 14.94 8.49 8.59 8.59 8.59 7.64 7.64 7.64CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 5.08 0.83 0.83 0.83 0.83 0.83 0.83 0.83 0.00 0.83 0.83 0.83 0.83 0.83 0.83N2 [kg/s] 0.00 0.00 0.00 0.25 0.00 0.40 0.40 0.40 0.40 0.40 0.40 0.40 0.40 0.00 0.40 0.40 0.40 0.40 0.40 0.40H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.01 0.01 0.01 0.01 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.07 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 23.88 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.65 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00sulfur [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.68 0.00 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.0 2.5 12.3 1.7 0.0 56.2 63.8 34.7 34.7 30.7 30.3 28.0 36.7 10.3 41.4 32.7 26.7 26.5 24.5 22.4ECH [MW] 488.1 7.7 0.5 0.6 0.0 356.5 395.1 395.1 395.0 395.0 394.8 394.8 395.2 0.4 388.6 388.6 388.6 387.9 387.9 387.9stream no. 227 230 231 232 233 234 235 236 238 239 241 242 244 247 250 252 253 254 255 256T [°C] 69.0 -31.1 145.0 145.0 15.0 30.0 20.0 20.0 20.0 15.0 20.1 980.0 -43.0 -70.5 30.0 30.9 318.1 318.1 480.0 400.0p [bar] 21.0 17.6 17.6 17.6 1.0 4.6 1.0 1.0 1.0 1.0 1.5 1.0 10.0 2.7 1.4 110.0 115.8 110.1 28.7 26.5m [kg/s] 50.0 8.1 5.5 2.5 25.5 25.5 0.1 18.0 5.7 37.7 2.5 40.2 10.8 11.8 11.8 34.3 32.7 32.7 0.0 0.0H [MW] -461.4 -41.5 -22.9 -10.4 -2.4 -2.2 -2.4 -0.3 0.0 -3.6 -12.2 -81.5 -96.6 -106.1 -105.1 -313.3 -475.8 -434.7 0.0 0.0O2 [kg/s] 0.00 0.00 0.00 0.00 5.89 5.89 0.00 0.44 5.45 8.73 0.00 0.79 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 38.04 3.80 2.61 1.19 0.00 0.00 0.00 0.00 0.00 0.00 1.19 2.03 10.71 11.75 11.76 34.22 0.00 0.00 0.00 0.00CO [kg/s] 0.26 0.25 0.18 0.08 0.00 0.00 0.00 0.00 0.00 0.00 0.08 0.00 0.01 0.00 0.00 0.01 0.00 0.00 0.00 0.00H2 [kg/s] 2.81 2.76 1.90 0.86 0.00 0.00 0.00 0.00 0.00 0.00 0.86 0.00 0.05 0.00 0.00 0.05 0.00 0.00 0.00 0.00H2O [kg/s] 7.64 0.00 0.00 0.00 0.16 0.16 0.15 0.01 0.00 0.24 0.00 8.54 0.00 0.00 0.00 0.00 32.74 32.74 0.00 0.00CH4 [kg/s] 0.83 0.83 0.57 0.26 0.00 0.00 0.00 0.00 0.00 0.00 0.26 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.40 0.40 0.27 0.12 19.41 19.41 0.00 17.52 0.25 28.74 0.12 28.86 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.01 0.01SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00sulfur [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 17.6 10.5 8.0 3.6 0.0 3.2 0.0 0.0 0.0 0.0 0.5 32.4 1.5 0.8 0.2 7.4 15.2 36.2 0.0 0.0ECH [MW] 387.9 369.4 254.0 115.4 0.1 0.1 0.0 0.3 0.6 0.2 115.4 1.4 10.3 5.3 5.3 20.8 1.6 1.6 0.2 0.2stream no. 257 259 260 261 262 263 264 265 266 269 270 271 272 273 274 275 276 277 280 281T [°C] 318.1 318.1 318.1 318.1 318.1 160.0 160.0 160.0 160.0 49.0 155.0 49.0 119.9 30.0 160.0 100.0 160.0 100.0 318.1 -30.0p [bar] 110.1 115.8 110.1 115.8 110.1 6.5 6.2 6.5 6.2 6.3 6.3 6.3 6.3 21.0 6.2 6.2 6.2 6.2 115.8 20.9m [kg/s] 5.3 3.2 3.2 12.1 12.1 6.1 6.1 2.4 2.4 15.4 15.4 58.1 58.1 7.6 1.1 1.1 0.4 0.4 5.3 0.1H [MW] -69.8 -47.0 -42.9 -175.5 -160.4 -94.0 -81.2 -36.7 -31.7 -242.4 -235.6 -924.2 -904.2 -121.2 -14.1 -16.6 -5.3 -6.2 -76.4 -1.2O2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.01 0.00 0.00 0.00 0.00 0.00 0.00CO [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 5.26 3.23 3.23 12.08 12.08 6.14 6.14 2.40 2.40 15.37 15.37 58.05 58.05 7.56 1.07 1.07 0.40 0.40 5.26 0.07CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00sulfur [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 5.8 1.5 3.6 5.6 13.4 0.7 5.0 0.3 2.0 0.1 1.7 0.6 4.4 0.0 0.9 0.0 0.3 0.0 2.4 0.0ECH [MW] 0.3 0.2 0.2 0.6 0.6 0.3 0.3 0.1 0.1 0.8 0.8 2.9 2.9 0.4 0.1 0.1 0.0 0.0 0.3 0.0stream no. 287 288 289 290 292 293 294 295 296 297 298 299 300 301 302 303 304 305 306 308T [°C] 41.2 105.0 41.3 35.2 90.0 18.6 145.0 30.1 -9.6 23.6 15.0 15.0 24.9 980.0 1011.7 980.0 30.0 20.0 20.0 30.0p [bar] 33.0 1.2 1.0 1.2 1.0 1.1 17.6 6.3 1.5 6.3 1.0 1.0 1.1 1.0 1.1 1.0 60.0 1.0 1.0 60.0m [kg/s] 26.5 36.5 1291.2 36.5 1270.0 1270.0 8.1 15.4 2.5 15.4 0.7 1270.0 37.7 27.8 27.8 68.1 0.4 0.4 1.3 0.2H [MW] -164.2 -485.1 -371.5 -577.6 -24.0 -116.5 -33.4 -243.7 -12.6 -244.1 -10.4 -121.2 -3.2 -56.4 -55.1 -137.8 0.0 0.0 0.0 0.0O2 [kg/s] 0.00 0.00 293.96 0.00 293.96 293.96 0.00 0.00 0.00 0.00 0.00 293.96 8.73 0.55 0.55 1.34 0.00 0.00 0.00 0.00CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 3.80 0.00 1.19 0.00 0.00 0.00 0.00 1.40 1.40 3.43 0.00 0.00 0.00 0.00CO [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.25 0.00 0.08 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 2.76 0.00 0.86 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 2.65 36.48 29.27 36.48 8.04 8.04 0.00 15.37 0.00 15.37 0.00 8.04 0.24 5.91 5.91 14.45 0.00 0.00 0.00 0.00CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.83 0.00 0.26 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.00 0.00 968.00 0.00 968.00 968.00 0.40 0.00 0.12 0.00 0.00 968.00 28.74 19.98 19.98 48.84 0.37 0.37 1.26 0.24H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 23.88 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00sulfur [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.1 21.3 3.4 0.1 11.0 3.9 11.6 0.0 0.4 0.0 0.0 0.0 0.3 22.4 23.7 54.8 0.1 0.0 0.0 0.1ECH [MW] 487.0 1.8 6.6 1.8 5.8 5.8 369.4 0.8 115.4 0.8 0.0 5.8 0.2 1.0 1.0 2.3 0.0 0.0 0.0 0.0

[MWel]W1 4.58 ASU air compressor (K21)W2 2.99 O2 compressor (K23)W3 0.57 ASU cold box (K22)W4 4.01 AGR (K19)W5 9.29 CO2 compressor (K20)W7 2.13 syngas expander (K25)W8 26.23 total electricity consumptionW9 4.91 drier fan (K28)W10 0.40 reformer air compressor (K30)W11 1.38 reformer recycle compressor (K31)W12 0.23 N2 compressor (K32)

295


Table C.5: FB-wood-3 flow stream data.

stream no. 201 205 207 208 209 211 213 214 215 216 217 218 219 220 221 222 223 224 225T [°C] 15.0 899.8 359.6 163.3 15.0 900.0 950.0 480.0 480.0 400.0 399.8 350.0 349.1 350.0 491.3 328.1 200.0 223.4 170.0p [bar] 1.0 1.0 37.3 37.3 33.3 31.6 29.6 28.7 28.7 28.5 26.5 26.4 26.4 26.9 24.1 24.1 23.6 21.3 21.0m [kg/s] 47.8 2.0 9.8 5.7 0.7 41.5 46.6 46.6 46.6 46.6 46.6 46.6 51.6 4.9 51.6 51.6 51.6 51.6 51.6H [MW] -507.1 -16.4 -125.6 0.7 -8.6 -305.7 -303.6 -346.7 -346.7 -353.7 -353.7 -357.9 -421.3 -63.4 -421.7 -438.1 -450.4 -450.4 -455.5O2 [kg/s] 0.00 0.00 0.00 5.45 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 20.20 20.20 20.20 20.20 20.20 20.20 20.20 20.20 0.00 37.07 37.07 37.07 39.61 39.61CO [kg/s] 0.00 0.00 0.00 0.00 0.00 4.06 12.64 12.64 12.64 12.64 12.64 12.64 12.64 0.00 1.90 1.90 1.90 0.28 0.28H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.38 1.60 1.60 1.60 1.60 1.60 1.60 1.60 0.00 2.37 2.37 2.37 2.49 2.49H2O [kg/s] 23.88 0.00 9.78 0.00 0.00 11.30 11.32 11.32 11.32 11.32 11.32 11.32 16.26 4.94 9.35 9.35 9.35 8.31 8.31CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 5.08 0.25 0.25 0.25 0.25 0.25 0.25 0.25 0.00 0.25 0.25 0.25 0.25 0.25N2 [kg/s] 0.00 0.00 0.00 0.25 0.00 0.40 0.63 0.63 0.63 0.63 0.63 0.63 0.63 0.00 0.63 0.63 0.63 0.63 0.63H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.01 0.01 0.01 0.01 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.07 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 23.88 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.65 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00sulfur [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.68 0.00 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.0 2.5 12.3 1.7 0.0 56.2 68.4 37.9 37.9 33.7 33.3 30.9 36.1 6.0 41.2 31.8 26.0 25.9 23.8ECH [MW] 488.1 7.7 0.5 0.6 0.0 356.5 329.0 329.0 328.8 328.8 328.7 328.7 328.9 0.2 321.6 321.6 321.6 320.8 320.8stream no. 226 227 230 231 233 234 235 236 238 239 244 247 250 252 253 254 255 256 257T [°C] 134.8 69.0 -31.3 145.0 15.0 30.0 20.0 20.0 20.0 20.0 -43.0 -70.5 30.0 30.9 318.1 318.1 480.0 399.8 318.1p [bar] 21.0 21.0 17.6 17.6 1.0 4.6 1.0 1.0 1.0 1.0 10.0 2.7 1.4 110.0 115.8 110.1 28.7 26.5 110.1m [kg/s] 51.6 51.6 7.5 7.5 48.4 48.4 0.3 34.2 10.8 5.1 11.2 12.2 12.2 35.7 34.3 34.3 0.0 0.0 5.5H [MW] -461.5 -483.7 -39.9 -32.8 -4.6 -4.1 -4.6 -0.5 -0.1 0.0 -100.6 -110.4 -109.5 -326.2 -498.7 -455.5 0.0 0.0 -73.2O2 [kg/s] 0.00 0.00 0.00 0.00 11.20 11.20 0.00 0.84 10.36 4.91 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 39.61 39.61 3.96 3.96 0.00 0.00 0.00 0.00 0.00 0.00 11.15 12.24 12.24 35.64 0.00 0.00 0.00 0.00 0.00CO [kg/s] 0.28 0.28 0.28 0.28 0.00 0.00 0.00 0.00 0.00 0.00 0.01 0.00 0.00 0.01 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 2.49 2.49 2.44 2.44 0.00 0.00 0.00 0.00 0.00 0.00 0.05 0.00 0.00 0.05 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 8.31 8.31 0.00 0.00 0.31 0.31 0.28 0.02 0.00 0.00 0.00 0.00 0.00 0.00 34.32 34.32 0.00 0.00 5.51CH4 [kg/s] 0.25 0.25 0.25 0.25 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.63 0.63 0.62 0.62 36.88 36.88 0.00 33.29 0.48 0.23 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.01 0.01 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00sulfur [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 21.9 16.5 9.2 10.2 0.0 6.1 0.0 0.0 0.0 0.0 1.6 0.8 0.2 7.7 15.9 38.0 0.0 0.0 6.1ECH [MW] 320.8 320.8 301.2 301.2 0.2 0.2 0.0 0.6 1.1 0.5 10.7 5.6 5.5 21.7 1.7 1.7 0.2 0.2 0.3stream no. 259 260 261 262 263 264 265 266 269 270 271 272 273 274 275 276 277 280 281T [°C] 318.1 318.1 318.1 318.1 160.0 160.0 160.0 160.0 49.0 155.0 49.0 119.6 30.0 160.0 100.0 160.0 100.0 318.1 -30.0p [bar] 115.8 110.1 115.8 110.1 6.5 6.2 6.5 6.2 6.3 6.3 6.3 6.3 21.0 6.2 6.2 6.2 6.2 115.8 20.9m [kg/s] 3.4 3.4 13.0 13.0 5.9 5.9 2.4 2.4 13.4 13.4 64.9 64.9 8.3 1.1 1.1 0.8 0.8 5.5 0.1H [MW] -49.2 -44.9 -188.9 -172.6 -91.0 -78.6 -37.3 -32.2 -211.8 -205.8 -1033.9 -1011.7 -132.0 -14.7 -17.3 -10.0 -11.8 -80.1 -1.1O2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.01 0.00 0.00 0.00 0.00 0.00 0.00CO [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 3.38 3.38 13.00 13.00 5.95 5.95 2.44 2.44 13.43 13.43 64.95 64.95 8.24 1.11 1.11 0.76 0.76 5.51 0.07CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00sulfur [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 1.6 3.7 6.0 14.4 0.7 4.8 0.3 2.0 0.1 1.5 0.6 4.9 0.0 0.9 0.0 0.6 0.0 2.6 0.0ECH [MW] 0.2 0.2 0.6 0.6 0.3 0.3 0.1 0.1 0.7 0.7 3.2 3.2 0.4 0.1 0.1 0.0 0.0 0.3 0.0stream no. 287 288 289 290 292 293 294 295 298 299 300 301 304 305 306 308T [°C] 41.2 105.0 41.3 35.2 90.0 18.6 145.0 29.7 15.0 15.0 798.3 20.0 30.0 20.0 20.0 30.0p [bar] 33.0 1.2 1.0 1.2 1.0 1.1 17.6 6.3 1.0 1.0 31.8 1.0 60.0 1.0 1.0 60.0m [kg/s] 26.5 36.5 1291.2 36.5 1270.0 1270.0 7.5 13.4 0.7 1270.0 5.1 5.7 0.4 0.4 2.7 0.2H [MW] -164.2 -485.1 -371.5 -577.6 -24.0 -116.5 -32.8 -212.9 -10.4 -121.2 4.1 0.0 0.0 0.0 0.0 0.0O2 [kg/s] 0.00 0.00 293.96 0.00 293.96 293.96 0.00 0.00 0.00 293.96 4.91 5.45 0.00 0.00 0.00 0.00CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 3.96 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.28 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 2.44 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 2.65 36.48 29.27 36.48 8.04 8.04 0.00 13.43 0.00 8.04 0.00 0.00 0.00 0.00 0.00 0.00CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.25 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.00 0.00 968.00 0.00 968.00 968.00 0.62 0.00 0.00 968.00 0.23 0.25 0.37 0.37 2.73 0.24H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 23.88 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00sulfur [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.1 21.3 3.4 0.1 11.0 3.9 10.2 0.0 0.0 0.0 3.5 0.0 0.1 0.0 0.0 0.1ECH [MW] 487.0 1.8 6.6 1.8 5.8 5.8 301.2 0.7 0.0 5.8 0.5 0.6 0.0 0.0 0.1 0.0

[MWel]W1 8.69 ASU air compressor (K21)W2 2.99 O2 compressor (K23)W3 1.08 ASU cold box (K22)W4 4.18 AGR (K19)W5 9.67 CO2 compressor (K20)W8 36.07 total electricity consumptionW9 4.91 drier fan (K28)W10 4.31 reformer air compressor (K30)W12 0.23 N2 compressor (K32)296

C.2 IGCC

C.2 IGCC

C.2.1 Simulation data from EF-IGCC-HTC-1

Table C.6: EF-IGCC-HTC-1 flow stream data from the gasification section.


stream no. 21 22 23 24 26 27 28 30 31 32 33 34 35 36 37 39 40 41T [°C] 166.9 161.8 160.6 133.3 38.0 142.7 253.2 25.0 150.1 241.2 71.1 165.0 35.0 254.7 160.1 1200.0 838.8 79.4p [bar] 30.7 30.7 30.7 30.7 30.5 43.0 43.0 1.1 1.0 1.0 1.0 7.0 7.0 43.0 43.0 39.0 39.0 25.0m [kg/s] 294.5 294.5 294.5 294.5 225.6 65.6 5.3 32.0 32.0 32.0 36.6 1.6 1.6 1.4 1.4 1.5 1.5 15.0H [MW] -2601.7 -2627.4 -2632.5 -2722.0 -1741.4 6.9 -69.5 -0.3 3.9 7.0 -60.4 -21.1 -25.2 -19.0 -22.1 4.9 2.8 -236.6O2 [kg/s] 0.00 0.00 0.00 0.00 0.00 62.72 0.00 0.79 0.79 0.79 0.79 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 191.38 191.38 191.38 191.38 191.26 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO [kg/s] 7.00 7.00 7.00 7.00 7.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 11.98 11.98 11.98 11.98 11.98 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 69.20 69.20 69.20 69.20 0.42 0.00 5.31 0.02 0.02 0.02 4.64 1.59 1.59 1.44 1.44 0.00 0.00 15.00CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 14.93 14.93 14.93 14.93 14.93 2.89 0.00 31.19 31.19 31.19 31.19 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/s] 0.05 0.05 0.05 0.05 0.05 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1.37 1.37 0.00EPH [MW] 155.3 146.5 144.8 116.4 90.3 19.9 6.1 0.1 0.8 2.0 1.3 1.3 0.0 1.5 0.2 3.3 1.7 0.5ECH [MW] 1537.6 1537.6 1537.6 1537.6 1534.5 6.9 0.3 0.6 0.6 0.6 0.8 0.1 0.1 0.1 0.1 5.7 5.7 0.7


stream no. 62 63 64 65 66 67 68 70 71 72 73 74 80 81 82T [°C] 144.4 120.8 128.8 32.9 108.1 72.4 120.0 166.9 253.2 248.3 160.0 256.8 25.0 130.6 122.0p [bar] 150.2 48.2 47.9 2.6 2.6 3.0 3.0 30.7 43.0 31.1 47.8 44.5 1.1 30.7 25.0m [kg/s] 272.2 130.4 130.4 246.4 246.4 15.0 15.0 9.4 116.9 304.0 24.1 24.1 32.0 294.5 15.0H [MW] -4181.8 -2030.0 -2024.9 -3940.6 -3851.1 -235.2 -232.2 -144.6 -1531.8 -2680.5 -368.8 -317.7 -0.3 -2728.1 -233.4O2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.79 0.00 0.00CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.03 0.00 191.41 0.00 0.00 0.00 191.38 0.00CO [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 7.00 0.00 0.00 0.00 7.00 0.00H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 11.98 0.00 0.00 0.00 11.98 0.00H2O [kg/s] 272.20 130.41 130.41 246.35 246.35 15.00 15.00 9.41 116.92 78.61 24.10 24.10 0.02 69.20 15.00CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 14.93 0.00 0.00 31.19 14.93 0.00H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.05 0.00 0.00 0.00 0.05 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 29.7 10.7 12.1 0.7 14.9 0.3 1.0 1.4 133.4 182.4 2.9 25.6 0.1 114.7 1.2ECH [MW] 13.6 6.5 6.5 12.3 12.3 0.7 0.7 0.5 5.8 1538.1 1.2 1.2 0.6 1537.6 0.7

297


Table C.7: EF-IGCC-HTC-1 flow stream data from the air separation, acid gas removaland gas turbine system.

stream no. 101 102 103 104 105 107 108 110 111 112 113 114 115 118 119 121 122 123T [°C] 15.0 30.0 20.0 20.0 20.0 20.0 30.0 30.0 20.0 20.0 15.0 15.3 170.0 114.4 20.0 231.1 101.9 15.0p [bar] 1.0 15.4 4.0 4.0 4.0 4.0 60.0 60.0 4.0 4.0 1.0 23.0 22.8 18.0 4.0 17.6 17.6 1.0m [kg/s] 207.9 206.9 65.6 0.0 65.6 69.2 69.2 57.2 156.5 124.5 7.1 7.1 7.1 124.5 32.0 128.8 52.6 1130.1H [MW] -19.8 -3.6 -0.4 0.0 -0.4 -0.4 -0.4 -0.3 -2.3 -1.8 -113.4 -113.4 -108.0 10.3 -0.5 -30.2 -183.2 -107.8O2 [kg/s] 48.12 48.12 62.75 0.03 62.72 0.00 0.00 0.00 5.09 4.05 0.00 0.00 0.00 4.05 1.04 4.05 0.00 261.58CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 18.93 0.00CO [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 6.82 0.00H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 11.78 0.00H2O [kg/s] 1.32 0.30 0.00 0.00 0.00 0.00 0.00 0.00 0.10 0.08 7.05 7.05 7.05 0.08 0.02 4.37 0.23 7.15CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 158.47 158.47 2.89 0.00 2.89 69.15 69.15 57.20 151.33 120.37 0.00 0.00 0.00 120.37 30.94 120.37 14.79 861.36H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00sulphur [kg/s]EPH [MW] 0.0 46.9 6.8 0.0 6.8 8.1 24.2 20.0 18.3 14.6 0.0 0.0 1.1 32.4 3.7 40.0 50.7 0.0ECH [MW] 0.9 1.1 6.9 0.0 6.9 1.8 1.8 1.5 2.7 2.2 0.4 0.4 0.4 2.2 0.6 2.2 1444.5 5.1

stream no. 130 131 133 134 135 136 137 141 143 145 151 203 204 205 206 207 209 211T [°C] 387.1 387.1 120.0 166.0 266.6 387.1 1350.6 605.4 79.3 79.3 20.0 134.0 49.0 145.4 145.4 49.0 49.0 49.0p [bar] 15.6 15.6 15.2 4.0 7.9 15.6 15.1 1.1 17.8 17.8 4.0 27.0 1.5 4.2 4.2 2.9 9.2 23.3m [kg/s] 927.5 85.2 85.2 10.8 34.1 72.5 1108.8 1226.3 15.0 2.8 0.0 52.6 0.3 10.0 10.0 77.8 43.2 51.8H [MW] 271.7 24.9 0.9 0.6 5.6 21.2 32.7 -1004.9 -236.6 -43.6 0.0 -176.3 -1.8 -132.1 -153.3 -692.5 -384.9 -462.5O2 [kg/s] 214.68 19.71 19.71 2.50 7.90 16.78 121.35 148.54 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 29.66 29.66 0.00 0.00 0.00 18.93 0.19 0.00 0.00 77.46 43.03 51.64CO [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 6.82 0.00 0.00 0.00 0.08 0.04 0.05H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 11.77 0.00 0.00 0.00 0.09 0.05 0.06H2O [kg/s] 5.87 0.54 0.54 0.07 0.22 0.46 115.72 116.46 15.00 2.76 0.00 0.23 0.01 9.98 9.98 0.08 0.04 0.05CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 706.92 64.90 64.90 8.24 26.02 55.27 842.08 931.62 0.00 0.00 0.02 14.79 0.00 0.00 0.00 0.06 0.03 0.04H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.05 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00sulphur [kg/s]EPH [MW] 341.7 31.4 20.6 1.6 8.3 26.7 1485.0 416.7 0.5 0.1 0.0 59.8 0.0 7.6 1.0 4.7 5.3 8.8ECH [MW] 4.2 0.4 0.4 0.0 0.2 0.3 17.6 17.6 0.7 0.1 0.0 1444.1 1.2 0.5 0.5 45.8 25.5 30.6

stream no. 214 215 216 218 226 227 228 229 230 231 231 233 234 235T [°C] 30.0 42.6 49.0 34.0 34.0 160.0 256.8 160.0 166.8 165.0 100.0 165.0 100.0 15.0p [bar] 110.0 23.1 1.5 27.2 27.2 47.8 44.5 7.4 7.2 7.0 7.0 7.0 7.0 1.0m [kg/s] 172.6 0.2 0.0 0.0 52.6 0.0 0.0 0.0 0.0 21.7 21.7 4.6 4.6 0.044H [MW] -1577.3 -2.7 -0.3 0.0 -197.6 -0.2 -0.2 -0.3 -0.2 -287.2 -338.1 -60.3 -71.0 0.0O2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 172.13 0.00 0.00 0.00 18.93 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO [kg/s] 0.17 0.00 0.00 0.00 6.82 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 0.20 0.00 0.00 0.00 11.77 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 0.00 0.17 0.02 0.00 0.23 0.01 0.01 0.02 0.02 21.73 21.73 4.56 4.56 0.00CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.13 0.00 0.00 0.00 14.79 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00sulphur [kg/s] 0.044EPH [MW] 36.7 0.0 0.0 0.0 56.0 0.0 0.0 0.0 0.0 18.1 1.0 3.8 0.2 0.0ECH [MW] 102.0 0.0 0.0 0.1 1444.1 0.0 0.0 0.0 0.0 1.1 1.1 0.2 0.2 0.8

298

C.2 IGCC

Table C.8: EF-IGCC-HTC-1 flow stream data from the steam cycle.

No. m T p EPH ECH type[kg/s] [°C] [bar] [MW] [MW]

301 1226.28 605.4 1.1 411.8 17.6 exhaust gas302 478.69 605.4 1.1 160.7 6.9 exhaust gas303 747.59 605.4 1.1 251.0 10.7 exhaust gas304 747.59 433.5 1.1 151.7 10.7 exhaust gas305 478.69 433.8 1.1 97.2 6.9 exhaust gas306 273.96 585.0 127.3 445.8 13.7 steam308 251.65 585.0 40.8 385.5 12.6 steam309 273.96 335.1 137.3 299.6 13.7 steam310 250.00 335.3 137.3 273.4 12.5 steam311 23.96 337.0 140.4 26.1 1.2 steam312 1226.28 433.6 1.1 248.9 17.6 exhaust gas313 1226.28 366.9 1.0 193.5 17.6 exhaust gas314 273.96 395.0 134.8 351.6 13.7 steam315 1226.28 348.9 1.0 179.0 17.6 exhaust gas316 1226.28 169.8 1.0 65.3 17.6 exhaust gas320 273.96 335.9 147.6 143.6 13.7 water321 0.13 120.3 2.0 0.1 0.0 steam325 7.25 326.1 7.0 7.1 0.4 steam326 23.96 335.9 147.6 13.1 1.2 water327 26.99 254.7 43.0 28.6 1.3 steam328 251.65 414.8 43.0 324.9 12.6 steam332 1226.28 169.8 1.0 65.3 17.6 water334 1.22 160.0 739.9 0.2 0.1 water341 8.47 165.1 7.0 7.0 0.4 steam343 244.40 32.9 0.1 32.1 12.2 steam (x=91.28%)344 465.99 169.8 1.0 24.8 6.7 exhaust gas345 490.51 169.8 1.0 26.1 7.1 exhaust gas346 269.78 169.8 1.0 14.4 3.9 exhaust gas347 465.99 155.0 1.0 22.2 6.7 exhaust gas348 273.96 122.0 151.2 22.4 13.7 water349 273.96 151.6 150.0 32.6 13.7 water350 490.51 139.9 1.0 20.7 7.1 exhaust gas351 127.35 130.0 47.9 10.4 6.4 water352 127.35 160.0 47.8 15.4 6.4 water353 269.78 158.2 1.0 13.1 3.9 exhaust gas354 20.64 120.3 7.5 1.4 1.0 water355 20.64 160.0 7.4 2.4 1.0 water356 250.00 335.9 147.6 131.0 12.5 water357 122.67 160.0 47.8 14.9 6.1 water358 19.42 160.0 7.4 2.3 1.0 water359 1226.28 149.7 1.0 55.8 17.6 exhaust gas360 273.96 145.0 150.2 30.1 13.7 water361 127.35 120.1 2.0 8.3 6.4 water362 20.64 120.1 2.0 1.3 1.0 water364 244.40 32.8 0.1 0.5 12.2 water365 244.40 120.0 2.5 15.9 12.2 water368 244.40 32.9 2.6 0.6 12.2 water371 273.96 120.1 2.0 17.9 13.7 water372 127.35 120.8 48.2 9.0 6.4 water373 148.35 15.0 1.0 0.0 7.4 water376 1.30 146.1 4.2 1.0 0.1 steam377 1226.28 102.2 1.0 37.5 17.6 exhaust gas378 148.35 118.8 2.5 9.5 7.4 water380 148.35 15.0 2.6 0.0 7.4 water381 11.28 120.1 2.0 0.7 0.6 water384 22.31 414.8 43.0 28.8 1.1 steam386 4.68 160.0 47.8 0.6 0.2 water387 11.28 120.2 4.4 0.7 0.6 water388 26.29 100.0 7.4 1.2 1.3 water389 1.59 33.0 7.4 0.0 0.1 water390 9.98 145.0 4.4 1.0 0.5 water392 1.44 160.2 43.0 0.2 0.1 water

299


Table C.9: EF-IGCC-HTC-1 mechanical and electrical work data.

No. [MW] component no. description

production21 1054.33 K35 gas turbine expander shaft work

16 -436.25 K33 gas turbine compressor shaft work

22 608.81 gas turbine system net electric output

44 392.83 K62 steam turbine electric output

48 2.71 K82 N2 expander (ASU)

50 6.58 K87 syngas expander

46 1010.93 plant gross electric output

47 770.10 plant net electric output

consumption1 4.40 K3 raw gas recycle compressor

3 16.75 K22 coal mill

4 0.27 cooling water pumps for compressor intercooling

5 0.77 cooling water pumps for condenser

6 2.00 miscellaneous

7 65.15 K37 ASU air compressor

8 22.14 K39 ASU N2 compressor

9 24.48 K40 ASU high pressure N2 compressor

10 20.65 K41 ASU O2 compressor

24 0.02 K44 saturator pump

25 0.03 K43 saturator pump

26 6.56 K38 ASU cold box

30 36.64 K26 CO2 compressor

32 34.94 K25 AGR

37 5.11 K68 feedwater pump



40 0.08 K71 condensate pump



300

C.2 IGCC

Table C.10: EF-IGCC-HTC-1 steam production and consumption.

HP steam

productionraw gas cooler [kg/s] 250.002

HRSG HP evaporator [kg/s] 23.957

consumptionsteam turbine [kg/s] 273.959

MP steam

productionraw gas coolers [kg/s] 93.692

gasifier membrane wall cooling [kg/s] 2.969

Claus plant [kg/s] 0.013

steam turbine extraction [kg/s] 26.995

consumptiongasifier (gasification agent) [kg/s] 5.306

shift reactor (reactant) [kg/s] 116.919

drier [kg/s] 1.444

LP steam


Claus plant [kg/s] 0.019

steam turbine extraction [kg/s] 8.469

consumptionAGR [kg/s] 21.728

ASU [kg/s] 4.564

drier [kg/s] 1.594

VP steam


consumptiondeaerator [kg/s] 1.300

syngas preheater [kg/s] 9.976

301


C.2.2 Simulation data from FB-IGCC-wood-1

Table C.11: FB-IGCC-wood-1 flow stream data from the gasification section, air separa-tion, acid gas removal and gas turbine system.

stream no. 201 205 207 208 209 211 213 214 215 216 217 218 219 220 221 222 223 224 225 226 227 230T [°C] 15.0 899.8 359.6 163.3 15.0 900.0 950.0 480.0 480.0 400.0 400.0 350.0 344.3 329.0 471.4 333.0 200.0 220.9 172.0 127.7 43.4 -31.1p [bar] 1.0 1.0 37.3 37.3 33.3 31.6 29.6 28.7 28.7 28.5 26.5 26.4 26.4 26.9 24.1 24.1 23.6 21.3 21.0 21.0 21.0 17.6m [kg/s] 47.8 2.0 9.8 5.7 0.7 41.5 41.5 41.5 41.4 41.4 41.4 41.4 50.0 8.5 50.0 50.0 50.0 50.0 50.0 50.0 50.0 8.1H [MW] -507.1 -16.4 -125.6 0.7 -8.6 -305.2 -239.8 -280.9 -280.9 -287.5 -287.5 -291.6 -401.6 -110.0 -401.9 -416.2 -429.5 -429.5 -434.3 -442.5 -465.3 -41.5O2 [kg/s] 0.00 0.00 0.00 5.41 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 20.16 20.16 20.16 20.16 20.16 20.16 20.16 20.16 0.00 35.69 35.69 35.69 38.02 38.02 38.02 38.02 3.80CO [kg/s] 0.00 0.00 0.00 0.00 0.00 4.06 11.62 11.62 11.62 11.62 11.62 11.62 11.62 0.00 1.74 1.74 1.74 0.26 0.26 0.26 0.26 0.26H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.38 2.00 2.00 2.00 2.00 2.00 2.00 2.00 0.00 2.71 2.71 2.71 2.81 2.81 2.81 2.81 2.77H2O [kg/s] 23.88 0.00 9.78 0.00 0.00 11.28 6.42 6.42 6.42 6.42 6.42 6.42 14.95 8.53 8.60 8.60 8.60 7.64 7.64 7.64 7.64 0.00CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 5.09 0.84 0.84 0.84 0.84 0.84 0.84 0.84 0.00 0.84 0.84 0.84 0.84 0.84 0.84 0.84 0.84N2 [kg/s] 0.00 0.00 0.00 0.25 0.00 0.40 0.40 0.40 0.40 0.40 0.40 0.40 0.40 0.00 0.40 0.40 0.40 0.40 0.40 0.40 0.40 0.39H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.01 0.01 0.01 0.01 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.07 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 23.88 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.65 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.68 0.00 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.0 2.5 12.3 1.7 0.0 56.1 63.8 34.7 34.7 30.7 30.3 28.0 36.5 10.1 41.2 33.0 26.7 26.5 24.6 22.1 17.2 10.5ECH [MW] 488.1 7.7 0.5 0.6 0.0 357.1 395.7 395.7 395.6 395.6 395.4 395.4 395.9 0.4 389.2 389.2 389.2 388.5 388.5 388.5 388.5 370.0stream no. 232 233 234 235 238 239 242 244 247 250 252 253 254 255 256 257 259 260 261 262 263 264T [°C] 145.0 15.0 30.0 20.0 20.0 15.0 980.0 -43.0 -70.5 30.0 30.9 323.0 324.3 480.0 400.0 324.3 323.0 324.3 323.0 324.3 162.0 160.0p [bar] 17.6 1.0 4.6 1.0 1.0 1.0 1.1 10.0 2.7 1.4 110.0 125.6 119.3 28.7 26.5 119.3 125.6 119.3 125.6 119.3 6.5 6.2m [kg/s] 2.5 25.3 25.3 0.1 5.7 37.6 40.2 10.8 11.7 11.8 34.3 34.0 34.0 0.0 0.0 5.5 3.4 3.4 11.8 11.8 6.4 6.4H [MW] -10.4 -2.4 -2.2 -2.4 0.0 -3.6 -81.3 -96.5 -106.0 -105.1 -313.1 -493.6 -452.5 0.0 0.0 -72.7 -48.8 -44.7 -171.7 -157.4 -98.0 -84.8O2 [kg/s] 0.00 5.85 5.85 0.00 5.41 8.71 0.79 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO2 [kg/s] 1.18 0.00 0.00 0.00 0.00 0.00 2.03 10.71 11.75 11.75 34.21 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO [kg/s] 0.08 0.00 0.00 0.00 0.00 0.00 0.00 0.01 0.00 0.00 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 0.86 0.00 0.00 0.00 0.00 0.00 0.00 0.05 0.00 0.00 0.05 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 0.00 0.16 0.16 0.15 0.00 0.24 8.52 0.00 0.00 0.00 0.00 34.04 34.04 0.00 0.00 5.47 3.36 3.36 11.84 11.84 6.41 6.41CH4 [kg/s] 0.26 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.12 19.25 19.25 0.00 0.25 28.70 28.82 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.01 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 3.6 0.0 3.2 0.0 0.0 0.0 32.5 1.5 0.8 0.2 7.4 16.3 37.6 0.0 0.0 6.0 1.6 3.7 5.7 13.1 0.8 5.2ECH [MW] 115.2 0.1 0.1 0.0 0.6 0.2 1.4 10.2 5.3 5.3 20.8 1.7 1.7 0.2 0.2 0.3 0.2 0.2 0.6 0.6 0.3 0.3stream no. 265 266 269 270 271 272 273 274 275 276 277 280 281 287 288 289 290 292 293 294 295 296T [°C] 162.0 161.8 28.8 155.0 28.8 100.0 30.0 160.0 100.0 160.0 100.0 323.0 -30.0 50.7 169.1 50.8 35.0 93.9 18.6 145.0 30.1 -9.6p [bar] 6.5 6.5 6.6 6.5 6.6 6.5 21.0 6.2 6.2 6.2 6.2 125.6 20.9 33.0 1.2 1.0 1.2 1.0 1.1 17.6 6.5 1.5m [kg/s] 2.3 2.3 15.4 15.4 60.0 75.4 7.6 1.1 1.1 0.4 0.4 5.5 0.1 26.5 32.3 1498.3 32.3 1477.1 1181.0 8.1 15.4 2.5H [MW] -35.4 -30.6 -243.9 -235.7 -951.5 -1172.9 -121.2 -14.1 -16.6 -5.2 -6.1 -79.3 -1.2 -163.7 -425.0 -820.5 -511.0 -470.3 -108.3 -33.4 -243.8 -12.6O2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 309.97 0.00 309.97 273.36 0.00 0.00 0.00CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 6.51 0.00 6.51 0.00 3.80 0.00 1.18CO [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.26 0.00 0.08H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 2.77 0.00 0.86H2O [kg/s] 2.31 2.31 15.38 15.38 60.00 75.38 7.57 1.07 1.07 0.39 0.39 5.47 0.07 2.65 32.28 59.78 32.28 38.55 7.48 0.00 15.38 0.00CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.84 0.00 0.26N2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1122.07 0.00 1122.07 900.16 0.39 0.00 0.12H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 23.88 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.3 1.9 0.0 1.7 0.1 3.4 0.0 0.9 0.0 0.3 0.0 2.6 0.0 0.1 20.1 10.3 0.1 17.0 3.7 11.6 0.0 0.4ECH [MW] 0.1 0.1 0.8 0.8 3.0 3.8 0.4 0.1 0.1 0.0 0.0 0.3 0.0 487.0 1.6 9.4 1.6 8.3 5.3 370.0 0.8 115.2stream no. 297 298 299 302 304 305 306 308 309 310 311 312 314 315 317 318 325 326 330 331 332 333T [°C] 23.6 15.0 15.0 145.0 154.0 20.0 230.0 108.2 29.0 980.0 980.0 1010.4 20.0 30.0 30.0 15.0 399.1 1230.4 548.8 170.3 274.2 399.1p [bar] 6.5 1.0 1.0 17.6 17.6 1.0 17.6 1.0 1.2 1.1 1.1 1.2 1.0 60.0 60.0 1.0 15.6 15.1 1.1 4.0 7.9 15.6m [kg/s] 15.4 0.7 1181.0 5.5 27.4 19.1 2.8 296.1 37.6 68.1 27.9 27.9 0.4 0.4 0.2 228.5 200.6 228.0 255.9 2.8 8.4 16.7H [MW] -244.2 -10.4 -112.7 -23.0 -56.9 -0.3 -36.3 -448.0 -3.1 -137.8 -56.5 -55.2 0.0 0.0 0.0 -21.8 61.4 0.0 -189.7 0.2 1.4 5.1O2 [kg/s] 0.00 0.00 273.36 0.00 0.44 0.44 0.00 36.61 8.71 1.34 0.55 0.55 0.00 0.00 0.00 52.90 46.44 29.36 35.81 0.64 1.94 3.87CO2 [kg/s] 0.00 0.00 0.00 2.62 2.62 0.00 0.00 6.51 0.00 3.43 1.41 1.41 0.00 0.00 0.00 0.00 0.00 4.48 4.48 0.00 0.00 0.00CO [kg/s] 0.00 0.00 0.00 0.18 0.18 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 0.00 0.00 0.00 1.90 1.90 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 15.38 0.00 7.48 0.00 2.78 0.01 2.77 31.07 0.24 14.44 5.92 5.92 0.00 0.00 0.00 1.45 1.27 22.37 22.55 0.02 0.05 0.11CH4 [kg/s] 0.00 0.00 0.00 0.58 0.58 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.00 0.00 900.16 0.27 18.90 18.63 0.00 221.91 28.70 48.84 20.02 20.02 0.37 0.37 0.24 174.19 152.93 171.83 193.09 2.12 6.38 12.76H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.0 0.0 0.0 8.0 14.8 0.0 2.9 11.7 0.4 55.1 22.6 23.8 0.0 0.1 0.1 0.0 75.5 272.3 74.2 0.4 2.1 6.3ECH [MW] 0.8 0.0 5.3 254.8 252.6 0.4 0.1 4.2 0.2 2.3 1.0 1.0 0.0 0.0 0.0 1.0 0.9 3.0 3.1 0.0 0.0 0.1

302

C.2 IGCC

Table C.12: FB-IGCC-wood-1 flow stream data from the steam cycle.


48 296.09 616 1.1 103.6 4.2 exhaust gas50 92.24 400 1.1 71.3 2.9 exhaust gas51 203.85 616 1.1 32.3 1.3 exhaust gas52 92.24 616 1.1 50.1 4.2 exhaust gas53 296.09 378 1.0 42.3 4.2 exhaust gas54 296.09 337 1.0 18.6 4.2 exhaust gas55 296.09 185 1.0 18.5 4.2 exhaust gas56 296.09 185 1.0 16.6 4.2 exhaust gas57 296.09 170 1.0 10.4 4.2 exhaust gas58 296.09 108 0.1 0.1 1.2 water60 21.09 15 1.0 0.0 1.1 water61 77.45 29 1.0 0.1 3.9 water62 77.45 29 6.6 0.1 3.9 water63 75.38 29 6.6 0.1 3.8 water64 2.07 29 6.6 0.0 0.1 water69 78.91 98 6.5 3.4 3.9 water72 78.91 159 6.5 9.1 3.9 water73 79.31 162 6.5 9.5 4.0 water78 11.12 162 6.5 1.3 0.6 water79 1.46 100 6.5 0.1 0.1 water80 1.46 100 6.2 0.1 0.1 water84 2.40 162 6.5 0.3 0.1 water85 66.57 162 6.5 8.0 3.3 water86 66.57 164 129.7 9.0 3.3 water87 66.57 323 125.6 31.9 3.3 water90 11.87 323 125.6 5.9 0.6 water91 0.57 162 6.5 0.1 0.0 water92 0.57 163 37.3 0.1 0.0 water93 0.38 162 6.5 0.0 0.0 water94 0.38 163 26.9 0.0 0.0 water97 2.40 161 6.2 2.0 0.1 steam

101 9.66 161 6.2 7.9 0.5 steam102 9.66 165 6.2 7.9 0.5 steam103 11.87 324 119.3 13.0 0.6 steam108 66.57 324 119.3 73.4 3.3 steam109 66.57 596 110.0 109.0 3.3 steam110 9.21 424 37.3 11.9 0.5 steam111 9.78 360 37.3 11.8 0.5 steam113 8.15 377 26.9 9.8 0.4 steam114 8.53 329 26.9 9.7 0.4 steam115 49.21 377 26.9 59.0 2.5 steam116 49.21 596 25.2 73.5 2.5 steam118 0.40 388 6.5 0.4 0.0 steam122 24.08 33 0.1 3.2 1.2 steam (x=92.42%)123 92.24 400 1.1 17.0 1.3 exhaust gas124 203.85 367 1.1 33.2 2.9 exhaust gas126 2.77 230 17.6 2.7 0.1 steam128 32.28 168 1.2 20.0 1.6 steam129 24.08 33 1.0 0.1 1.2 water130 32.28 35 1.2 0.1 1.6 water131 2.11 536 17.6 2.9 0.1 steam132 0.67 162 6.5 0.1 0.0 water133 0.7 162 17.6 0.1 0.0 water

303


Table C.13: FB-IGCC-wood-1 mechanical and electrical work data.


production194.433 K36 gas turbine expander shaft work

-90.859 K34 gas turbine compressor shaft work

20 102.020 K37 gas turbine system net electric output


7 2.127 K25 syngas expander

180.398 plant gross electric output


consumption1 4.538 K21 ASU air compressor



4 4.012 K19 AGR

5 9.303 K20 CO2 compressor

9 4.569 K29 drier air fan

10 7.582 K32 ASU N2 compressor


12 0.589 K30 reformer air compressor

13 1.353 K31 reformer recycle compressor


31 0.000 K53 pump




38 0.005 K56 pump

39 0.002 K57 pump

40 0.002 K58 pump



304

C.2 IGCC

Table C.14: FB-IGCC-wood-1 steam production and consumption.

HP steam


HRSG HP evaporator [kg/s] 11.869

consumptionsteam turbine [kg/s] 66.572

MP steam

productionsteam turbine extraction [kg/s] 21.09

consumptiongasifier (gasification agent) [kg/s] 9.78

shift reactor (reactant) [kg/s] 8.53

gas turbine injection [kg/s] 2.77

LP steam


HRSG LP evaporator [kg/s] 2.400

consumptionAGR [kg/s] 1.068

ASU [kg/s] 0.394

steam turbine [kg/s] 9.660

VP steam

productionsteam turbine extraction [kg/s] 32.677

consumptiondeaerator [kg/s] 0.402

drier [kg/s] 32.275

305


289

308

287

201

315

317

234314

293 299

305

276 277

238

235

233292

208

288 290

211

254

280257

253

214

255, 256

215

216

217

207

205 298

209

drier

feedingsystem

coldbox

exhaustgas

exhaustgas fromHRSG

exhaustgas todrier

wood

condensate

air

air

N2

N2

O2

N2

gasifier bed 1

bed 2

ash

sulfursand

dolomite

water

W1

W3W2W11

W9

solid fuel

steam

syngas

flue gas

electricity

CO2

liquid water

air, O , N2 2

K5 K6K23

K73

K27 K28

K22 K21K1

K33

K3

K7

K29 K8

A

A

318

333332331

325330

326exhaust

gasto HRSG

air

W20

GK34 K36K35

K37

B

B48330

51 123

103

52 53 55 57

97 78

72

85

93

9192

69

79 80

54 56

9087

86116

109 108

115

114

113110

111

118

128

122

129 6261

60

130

102

12450

HPHP

W30

G

K39

K48

K66K41 K43K42 K44

K45

K40

58

W36

W37

W38

W34

W31

K47

K62

K63K69K55

K54

K53

K67

K59

K56

73K60

C

C

D

D

E

E

F

F

Figure C.2: Flowsheet of the IGCC with fluidized bed gasifier without carbon capture,FB-IGCC-wood-0.

306

C.2 IGCC

C.2.3 Simulation data from FB-IGCC-wood-0

Table C.15: FB-IGCC-wood-0 flow stream data from the gasification section, air separa-tion, acid gas removal and gas turbine system.

stream no. 201 205 207 208 209 211 214 215 216 217 233 234 235 238 253 254 255 256T [°C] 15.0 899.8 359.6 163.3 15.0 900.0 480.0 480.0 400.0 399.8 15.0 30.0 20.0 20.0 323.0 324.3 480.0 399.8p [bar] 1.0 1.0 37.3 37.3 33.3 31.6 30.1 30.1 29.9 27.9 1.0 4.6 1.0 1.0 125.6 119.3 30.1 27.9m [kg/s] 47.8 2.0 2.7 4.8 0.7 33.4 33.4 33.4 33.4 33.4 21.3 21.3 0.1 4.8 23.0 23.0 0.0 0.0H [MW] -507.1 -16.4 -34.1 0.6 -8.6 -214.0 -241.8 -241.8 -246.6 -246.6 -2.0 -1.8 -2.0 0.0 -334.2 -306.3 0.0 0.0O2 [kg/s] 0.00 0.00 0.00 4.57 0.00 0.00 0.00 0.00 0.00 0.00 4.94 4.94 0.00 4.57 0.00 0.00 0.00 0.00CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 17.62 17.62 17.62 17.62 17.62 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00CO [kg/s] 0.00 0.00 0.00 0.00 0.00 5.01 5.01 5.01 5.01 5.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.22 0.22 0.22 0.22 0.22 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 23.88 0.00 2.65 0.00 0.00 4.68 4.68 4.68 4.68 4.68 0.14 0.14 0.13 0.00 23.04 23.04 0.00 0.00CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 5.48 5.48 5.48 5.48 5.48 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.00 0.00 0.00 0.21 0.00 0.36 0.36 0.36 0.36 0.36 16.26 16.26 0.00 0.21 0.00 0.00 0.00 0.00H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.01 0.01 0.01 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.01 0.01SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.07 0.07 0.07 0.07 0.07 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 23.88 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.65 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.68 0.00 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 0.0 2.5 3.3 1.4 0.0 41.6 22.0 22.0 19.2 18.9 0.0 2.7 0.0 0.0 11.0 25.4 0.0 0.0ECH [MW] 488.1 7.7 0.1 0.5 0.0 366.9 366.9 366.8 366.8 366.6 0.1 0.1 0.0 0.5 1.2 1.2 0.2 0.2stream no. 257 276 277 280 287 288 289 290 292 293 298 299 305 308 314 315 317 318T [°C] 324.3 160.0 100.0 323.0 47.9 145.1 48.1 35.0 88.8 18.6 15.0 15.0 20.0 85.1 20.0 30.0 30.0 15.0p [bar] 119.3 6.2 6.2 125.6 33.0 1.2 1.0 1.2 1.0 1.1 1.0 1.0 1.0 1.0 1.0 60.0 60.0 1.0m [kg/s] 4.0 0.3 0.3 4.0 26.5 32.0 1549.1 32.0 1527.8 1151.0 0.7 1151.0 16.1 376.8 0.4 0.4 0.2 343.4H [MW] -52.8 -4.4 -5.2 -57.5 -163.9 -423.3 -994.7 -507.1 -646.6 -105.6 -10.4 -109.8 -0.2 -624.9 0.0 0.0 0.0 -32.8O2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 319.23 0.00 319.23 266.42 0.00 266.42 0.37 52.81 0.00 0.00 0.00 79.49CO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 40.73 0.00 40.73 0.00 0.00 0.00 0.00 40.73 0.00 0.00 0.00 0.00CO [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 3.97 0.33 0.33 3.97 2.65 32.03 49.70 32.03 28.47 7.29 0.00 7.29 0.01 21.19 0.00 0.00 0.00 2.17CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 1139.41 0.00 1139.41 877.30 0.00 877.30 15.68 262.12 0.37 0.37 0.24 261.75H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 23.88 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.68 0.00 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 4.4 0.3 0.0 1.9 0.1 19.4 7.6 0.1 14.1 3.6 0.0 0.0 0.0 6.9 0.0 0.1 0.1 0.0ECH [MW] 0.2 0.0 0.0 0.2 487.0 1.6 15.7 1.6 14.6 5.2 0.0 5.2 0.3 13.2 0.0 0.0 0.0 1.6stream no. 325 326 330 331 332 333T [°C] 399.1 1229.6 559.6 170.3 274.2 399.1p [bar] 15.6 15.1 1.1 4.0 7.9 15.6m [kg/s] 302.3 335.8 376.8 4.1 12.3 24.6H [MW] 92.5 -158.6 -424.8 0.3 2.1 7.5O2 [kg/s] 69.98 43.31 52.82 0.95 2.85 5.70CO2 [kg/s] 0.00 40.73 40.73 0.00 0.00 0.00CO [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00H2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00H2O [kg/s] 1.91 20.93 21.19 0.03 0.08 0.16CH4 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00N2 [kg/s] 230.44 230.80 262.11 3.13 9.39 18.79H2S [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00SO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00tar [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00fuel DM [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00ash [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00SiO2 [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00dolomite [kg/s] 0.00 0.00 0.00 0.00 0.00 0.00EPH [MW] 113.7 383.9 106.1 0.6 3.1 9.3ECH [MW] 1.4 13.3 13.2 0.0 0.1 0.1

307


Table C.16: FB-IGCC-wood-0 flow stream data from the steam cycle.


48 376.85 560 1.1 105.1 13.2 exhaust gas50 128.15 400 1.1 69.4 8.7 exhaust gas51 248.70 560 1.1 35.8 4.5 exhaust gas52 128.15 560 1.1 64.5 13.2 exhaust gas53 376.85 411 1.0 47.2 13.2 exhaust gas54 376.85 339 1.0 27.1 13.2 exhaust gas55 376.85 240 1.0 26.3 13.2 exhaust gas56 376.85 236 1.0 15.9 13.2 exhaust gas57 376.85 172 1.0 6.0 13.2 exhaust gas58 376.85 85 0.1 0.1 1.5 water60 2.66 15 1.0 0.0 0.1 water61 65.10 33 1.0 0.2 3.3 water62 65.10 33 6.6 0.2 3.3 water69 65.43 34 6.5 0.2 3.3 water72 65.43 160 6.5 7.7 3.3 water73 65.66 162 6.5 7.9 3.3 water78 12.59 162 6.5 1.5 0.6 water79 0.33 100 6.5 0.0 0.0 water80 0.33 100 6.2 0.0 0.0 water85 52.98 162 6.5 6.4 2.6 water86 52.98 164 129.7 7.1 2.6 water87 52.98 323 125.6 25.4 2.6 water90 25.96 323 125.6 12.9 1.3 water91 0.04 162 6.5 0.0 0.0 water92 0.04 163 37.3 0.0 0.0 water93 0.05 162 6.5 0.0 0.0 water97 12.59 161 6.2 10.3 0.6 steam

102 12.59 220 6.2 10.9 0.6 steam103 25.96 324 119.3 28.5 1.3 steam108 52.98 324 119.3 58.3 2.6 steam109 52.98 540 110.0 81.8 2.6 steam110 2.61 375 37.3 3.2 0.1 steam111 2.65 360 37.3 3.2 0.1 steam113 0.28 336 6.2 0.3 0.0 steam114 0.33 161 6.2 0.3 0.0 steam115 50.36 331 26.9 57.4 2.5 steam116 50.36 540 25.2 71.0 2.5 steam118 0.23 342 6.5 0.2 0.0 steam122 30.42 33 0.1 4.0 1.5 steam (x=91.0%)123 128.15 400 1.1 21.0 4.5 exhaust gas128 32.02 145 1.2 19.4 1.6 steam129 30.42 33 1.0 0.1 1.5 water130 32.02 35 1.2 0.1 1.6 water

308

C.2 IGCC

Table C.17: FB-IGCC-wood-0 mechanical and electrical work data.


production273.356 K36 gas turbine expander shaft work

-136.585 K34 gas turbine compressor shaft work

20 134.719 K37 gas turbine system net electric output


206.995 plant gross electric output


consumption1 3.833 K21 ASU air compressor



9 4.453 K29 drier air fan






38 0.000 K56 pump



309


C.2.

4Co

stda

tafro

mEF

-IGCC

-HT

C-1

Tabl

eC.

18:E

F-IG

CC

-HT

C-1

equi

pmen

tlis

tw

ith

inve

stm

ent

cost

s.A

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the

heat

exch

ange

rco

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com

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nX

d/n

CB

M/n

[M€

]

CB

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[M€

]

spec

ifica

tion

sco

stfu

ncti

on

EF

gasi

fier

K1

100%

244

.21

kgfe

ed/s

15.6

2m

3/s

gas

135.

8927

1.78

incl

udin

gcy

clon

e,fil

ter

EF

G-1

=14

2.9

M€,

EF

G-2

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K3

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235

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mill

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K19

,K22

100%

246

.53

kg/s

87.3

417

4.67

EF

F-1

AG

RK

2510

0%2

5.54

7km

ol/s

2.17

3km

olC

O2/s

0.68

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55.7

111

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2810

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0.04

4kg

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2.33

2.33

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-1

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K33

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221

8.1

MW

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1016

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863.

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1.69

3.39

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K35

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252

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shaf

t15

.63

31.2

6G

EX

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es

K36

100%

230

4.4

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el25

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50.8

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ASU

air

com

pres

sor

K37

100%

232

.58

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box

K38

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231

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kgO

2/s

35.8

771

.73

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ASU

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N2

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K39

100%

211

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MW

el5.

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N2

com

pres

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K40

100%

212

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el6.

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310

C.2 IGCC

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HR

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220

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feed

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143

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20.

771

kg/s

12.7

825

.55

SH-1

acce

ssor

yel

ectr

icpl

ant

100%

177

0.1

MW

el11

0.14

110.

14A

EP

-1

mis

c.sy

ngas

trea

tmen

t10

0%2

5.54

7km

ol/s

4.51

9.02

MST

-1

cool

ing

wat

ersy

stem

100%

172

4.8

MW

th29

.82

29.8

2C

WS-

1

heat

exch

ange

rs,t

ot90

.87

CB

M,

tot

1257

.1

offsi

teco

st

land

2250

MW

HH

V0.

95O

SC-3

anci

llary

build

ings

2250

MW

HH

V15

.58

OSC

-2

site

deve

lopm

ent

2250

MW

HH

V15

.50

OSC

-1

311


com

pone

ntno

.f d

nX

d/n

CB

M/n

[M€

]

CB

M

[M€

]

spec

ifica

tion

sco

stfu

ncti

on

OC

,to

tal

32.0

3

CB

M+

OC

1289

.1

Tabl

eC.

19:E

F-IG

CC

-HT

C-1

heat

exch

ange

rlis

tw

ith

inve

stm

ent

cost

.

com

pone

ntno

.f d

nQ

d/n

[MW

]

Ad/

n

[m2

]

f M,1

for

HX

-1,

HX

-2

f M,2 for

HX

-3

CB

M/n

[M€

]

CB

M

[M€

]

cost

func

tion

syng

asco

oler

K4

100%

213

6.7

2415

3.73

1.55

3.11

HX

-1,H

X-2

,HX

-4

syng

asco

oler

K5

100%

240

.616

132.

731.

042.

08H

X-1

,HX

-2,H

X-4

syng

asco

oler

K7

100%

216

.397

02.

730.

761.

52H

X-1

,HX

-2,H

X-4

syng

asco

oler

K10

100%

10.

00

1.81

0.00

0.00

HX

-1,H

X-2

,HX

-4

syng

asco

oler

K11

100%

233

.244

211.

811.

643.

29H

X-1

,HX

-2,H

X-4

syng

asco

oler

K13

100%

212

.716

751.

810.

911.

81H

X-1

,HX

-2,H

X-4

syng

asco

oler

K14

100%

212

.725

091.

811.

162.

32H

X-1

,HX

-2,H

X-4

syng

asco

oler

K15

100%

213

.322

241.

811.

092.

17H

X-1

,HX

-2,H

X-4

drie

rH

XK

2012

0%2

2.5

265

1.00

0.11

0.22

HX

-1,H

X-2

drie

rH

XK

2112

0%2

1.8

236

1.00

0.10

0.21

HX

-1,H

X-2

syng

asco

oler

K23

100%

22.

535

21.

810.

350.

70H

X-1

,HX

-2,H

X-4

syng

asco

oler

K24

100%

244

.733

931.

811.

402.

79H

X-1

,HX

-2,H

X-4

syng

aspr

ehea

ter

K27

100%

210

.612

091.

000.

280.

56H

X-1

,HX

-2

wat

erpr

ehea

ter

K45

100%

22.

733

11.

000.

120.

25H

X-1

,HX

-2

N2

preh

eate

rK

4610

0%2

9.4

1629

1.00

0.33

0.67

HX

-1,H

X-2

HP

supe

rhea

ter

K47

100%

277

.890

795

1.81

2.0

7.81

15.6

2H

X-2

,HX

-3

rehe

ater

K48

100%

249

.884

110

1.81

2.0

7.38

14.7

6H

X-2

,HX

-3

HP

supe

rhea

ter

K49

100%

248

.145

671

1.81

2.0

4.72

9.45

HX

-2,H

X-3

HP

evap

orat

orK

5010

0%2

12.9

6344

1.00

1.0

0.99

1.97

HX

-2,H

X-3

HP

econ

omiz

erK

5110

0%2

125.

016

1764

1.00

1.0

7.87

15.7

4H

X-2

,HX

-3

HP

econ

omiz

erK

5710

0%2

3.8

5619

1.00

1.0

0.66

1.32

HX

-2,H

X-3

312

C.2 IGCC

com

pone

ntno

.f d

nQ

d/n

[MW

]

Ad/

n

[m2

]

f M,1

for

HX

-1,

HX

-2

f M,2 for

HX

-3

CB

M/n

[M€

]

CB

M

[M€

]

cost

func

tion

MP

econ

omiz

erK

5810

0%2

8.2

1671

11.

001.

01.

462.

93H

X-2

,HX

-3

LP

econ

omiz

erK

5910

0%2

1.7

1675

1.00

1.0

0.27

0.55

HX

-2,H

X-3

FW

preh

eate

rK

6010

0%2

32.3

1190

31.

001.

01.

142.

28H

X-2

,HX

-3

syng

asco

oler

K80

100%

225

.738

591.

811.

513.

02H

X-1

,HX

-2,H

X-4

syng

asco

oler

K81

100%

21.

524

51.

810.

280.

56H

X-1

,HX

-2,H

X-4

drie

rH

XK

8310

0%2

1.5

586

1.00

0.17

0.35

HX

-1,H

X-2

syng

asco

oler

K88

100%

21.

630

81.

810.

320.

65H

X-1

,HX

-2,H

X-4

90.8

7

313


C.2.

5Co

stda

tafro

mFB

-IGCC

-woo

d-1

Tabl

eC.

20:F

B-IG

CC

-woo

d-1

equi

pmen

tlis

tw

ith

inve

stm

ent

cost

s.A

brea

kdow

nof

the

heat

exch

ange

rco

sts

ispr

ovid

edin

Tabl

eC

.21.

com

pone

ntno

.f d

nX

d/n

CB

M/n

[M€

]

CB

M

[M€

]

spec

ifica

tion

sco

stfu

ncti

on

belt

drie

rK

112

0%3

8.49

2kg

ev/s

5.43

16.2

9B

D-1

=8.

03M

€,

BD

-2=

2.83

M€

FB

gasi

fier

K5

100%

211

.94

kgD

M/s

2.72

1m

3/s

15.6

231

.23

FB

G-1

=8.

52M

€,

FB

G-2

=9.

04M

€,

FB

G-3

=25

.57

M€

,

FB

G-4

=12

.2M

€,

FB

G-5

=22

.8M

€

cycl

one

100%

22.

721

m3/s

0.61

1.21

CY

C-1

,C

YC

-2

hot

gas

filte

r10

0%2

2.72

m3/s

8.19

16.3

7C

F-1

=8.

19M

€,

CF

-2=

11.6

9M

€,

CF

-3=

4.68

M€

stea

mre

form

erK

610

0%2

0.80

9kg

H2/s

15.2

630

.51

MSR

-1=

13.0

9M

€,

MSR

-2=

17.4

2M

€

wat

erga

ssh

iftK

1110

0%2

1.38

0km

ol/s

0.20

7km

olC

O/s

0.73

1.46

clea

nsh

iftW

GS-

2

AG

RK

1910

0%2

1.16

8km

ol/s

0.43

2km

olC

O/s

20.1

440

.28

AG

R-3

CO

2co

mpr

esso

rsK

2010

0%4

2.32

6M

Wel

2.93

11.7

1C

MP

-6

air

com

pres

sor

K21

100%

22.

269

MW

el1.

993.

99C

MP

-3

ASU

cold

box

K22

100%

22.

703

kgO

2/s

13.0

026

.00

low

pres

sure

ASU

-2

O2

com

pres

sors

K23

100%

21.

484

MW

el2.

014.

02C

MP

-5

syng

asex

pand

erK

2510

0%2

1.06

3M

Wel

0.70

1.40

SGE

-1

drie

rfa

nK

2811

0%3

356.

3m

3/s

0.16

0.47

AF

-1

desu

lphu

riza

tion

K29

100%

20.

006

kgS

/s0.

330.

65Zn

O-b

edH

DS-

1

314

C.2 IGCC

com

pone

ntno

.f d

nX

d/n

CB

M/n

[M€

]

CB

M

[M€

]

spec

ifica

tion

sco

stfu

ncti

on

air

com

pres

sor

K30

100%

20.

294

MW

el0.

511.

02C

MP

-3

recy

cle

com

pres

sor

K31

100%

20.

677

MW

el1.

032.

07C

MP

-4

N2

com

pres

sor

K32

100%

23.

791

MW

el2.

815.

62C

MP

-3

feed

ing

syst

emK

3310

0%2

13.2

7kg

/s1.

082.

15F

FB

-1,

FF

B-2

GT

com

pres

sor

K34

100%

245

.43

MW

shaf

t2.

264.

53G

AC

-2

GT

com

bust

orK

3510

0%2

152.

0M

WH

HV

0.49

0.98

GC

C-2

GT

expa

nder

K36

100%

297

.22

MW

shaf

t4.

508.

99G

EX

-2

GT

gene

rato

r,

auxi

liari

es

K37

100%

251

.01

MW

el1.

392.

77G

GE

-2

cond

ense

rK

4710

0%1

53.9

3M

Wth

0.97

0.97

CN

D-1

stea

mtu

rbin

eK

4810

0%1

76.2

5kW

el11

.24

11.2

4ST

C-1

feed

wat

erpu

mp

K53

100%

20.

055

kWel

0.00

0.00

P-1

feed

wat

erpu

mp

K54

100%

229

.84

kWel

0.03

0.05

P-1

cond

ensa

tepu

mp

K55

100%

22.

390

kWel

0.01

0.02

P-1

feed

wat

erpu

mp

K56

100%

22.

338

kWel

0.02

0.03

P-1

feed

wat

erpu

mp

K57

100%

21.

050

kWel

0.01

0.03

P-1

feed

wat

erpu

mp

K58

100%

21.

060

kWel

0.01

0.02

P-1

feed

wat

erpu

mp

K59

100%

267

9.7

kWel

0.41

0.82

P-1

N2

com

pres

sor

K73

100%

20.

122

MW

el0.

280.

56C

MP

-3

HR

SGdu

ct,c

asin

g,st

ack

100%

288

.44

MW

th

148

kgga

s/s

1.03

2.06

(HR

S-1a

+H

RS-

1b),

HR

S-2

HR

SGdr

ums

100%

238

.12

kgst

eam

/s1.

122.

24H

RD

-1

feed

wat

ersy

stem

100%

179

.31

kgF

W/s

9.11

9.11

FW

S-1

woo

dha

ndlin

g10

0%1

47.7

7kg

raw

/s

26.5

4kg

dri

ed/s

6.00

6.00

conv

eyor

,dri

edw

ood

chip

sst

orag

e,ir

onre

mov

al

(WC

-1,

WC

-2)+

(WS-

1,

WS-

2)+

IR-1

ash

hand

ling

100%

21.

011

kg/s

7.33

14.6

5A

H-1

acce

ssor

yel

ectr

icpl

ant

100%

114

2.8

MW

el40

.07

40.0

7A

EP

-1

mis

c.sy

ngas

trea

tmen

t10

0%2

1.38

0km

ol/s

1.70

3.41

MST

-1

315


com

pone

ntno

.f d

nX

d/n

CB

M/n

[M€

]

CB

M

[M€

]

spec

ifica

tion

sco

stfu

ncti

on

cool

ing

wat

ersy

stem

100%

179

.68

MW

th6.

366.

36C

WS-

1

heat

exch

ange

rs,t

ot32

.29

CB

M,

tot

343.

67

offsi

teco

st

land

467

MW

HH

V0.

32O

SC-3

anci

llary

build

ings

467M

WH

HV

5.61

OSC

-2

site

deve

lopm

ent

467M

WH

HV

5.58

OSC

-1

OC

,to

tal

11.5

0

CB

M+

OC

355.

17

316

C.2 IGCC

Tabl

eC.

21:F

B-IG

CC

-woo

d-1

heat

exch

ange

rlis

tw

ith

inve

stm

ent

cost

.

com

pone

ntno

.f d

nQ

d/n

[MW

]

Ad/

n

[m2

]

f M,1

for

HX

-1,

HX

-2

f M,2 for

HX

-3

CB

M/n

[M€

]

CB

M

[M€

]

cost

func

tion

syng

asco

oler

K7

100%

220

.56

303

3.73

0.43

20.

865

HX

-1,H

X-2

,HX

-4

syng

asco

oler

K8

100%

23.

3014

83.

730.

286

0.57

2H

X-1

,HX

-2,H

X-4

syng

asco

oler

K9

100%

22.

0321

51.

810.

263

0.52

6H

X-1

,HX

-2,H

X-4

syng

asco

oler

K12

100%

27.

1670

21.

810.

535

1.07

0H

X-1

,HX

-2,H

X-4

syng

asco

oler

K13

100%

26.

6537

31.

810.

361

0.72

3H

X-1

,HX

-2,H

X-4

syng

asco

oler

K15

100%

22.

4043

51.

810.

396

0.79

2H

X-1

,HX

-2,H

X-4

syng

asco

oler

K16

100%

24.

1044

01.

810.

399

0.79

8H

X-1

,HX

-2,H

X-4

syng

asco

oler

K17

100%

211

.41

2702

1.81

1.21

22.

425

HX

-1,H

X-2

,HX

-4

syng

aspr

ehea

ter

K24

100%

24.

0514

341

0.30

90.

619

HX

-1,H

X-2

syng

aspr

ehea

ter

K26

100%

20.

2110

91

0.06

80.

137

HX

-1,H

X-2

drie

rH

XK

2711

5%3

32.9

836

011

0.54

51.

635

HX

-1,H

X-2

rehe

ater

K39

100%

212

.11

1879

01.

812.

02.

475

4.94

9H

X-2

,HX

-3

HP

supe

rhea

ter

K40

100%

230

.67

3392

71.

812.

03.

803

7.60

6H

X-2

,HX

-3

HP

evap

orat

orK

4110

0%2

7.01

2367

11.

00.

472

0.94

3H

X-2

,HX

-3

HP

econ

omiz

erK

4210

0%2

25.8

729

837

11.

02.

243

4.48

6H

X-2

,HX

-3

LP

supe

rhea

ter

K43

100%

20.

0578

1.81

2.0

0.04

90.

099

HX

-2,H

X-3

LP

evap

orat

orK

4410

0%2

2.49

1733

11.

00.

374

0.74

8H

X-2

,HX

-3

LP

econ

omiz

erK

4510

0%2

10.2

319

690

11.

01.

651

3.30

2H

X-2

,HX

-3

32.2

9

317


C.2.6 Investment cost summary for the IGCC plants

Table C.22: Investment cost summary for IGCC plants.

EF EF

HTC-1

EF

TOR-1

FB

wood-1

FB

WP-1

FB

wood-0

feedstock bit. coal SR FR SR FR SR

coal mill, drier, lock hopper [M€] 172.3 187.4 211.0 21.1 2.3 21.3

gasifier [M€] 255.4 277.0 316.2 48.8 48.9 45.7

reformer [M€] — — — 33.6 33.5 —

syngas coolers [M€] 22.8 23.4 15.5 7.8 7.8 3.1

AGR [M€] 162.7 120.2 121.0 42.4 42.4 0.7

CO2 compressor [M€] 28.7 29.3 29.1 11.7 11.7 —

ASU [M€] 126.0 122.8 123.0 39.6 40.3 31.1

GT system [M€] 103.7 101.7 95.9 17.3 17.2 23.5

HRSG [M€] 81.0 84.6 85.3 26.4 26.3 19.4

steam turbine [M€] 46.2 49.1 50.8 11.2 13.0 10.7

condenser [M€] 7.1 7.7 8.2 1.0 2.1 1.2

feedwater pumps [M€] 2.7 2.7 2.6 1.0 1.0 0.8

feedwater system [M€] 34.9 35.5 34.3 9.1 9.1 7.8

cooling water system [M€] 28.4 29.8 31.1 6.4 10.0 5.9

coal handling [M€] 32.6 35.7 40.5 6.0 1.9 6.0

slag handling [M€] 96.7 25.6 34.1 14.7 14.7 14.7

accessory electric plant [M€] 109.4 110.1 107.5 40.1 42.9 48.1

misc. syngas treatment [M€] 12.0 12.0 11.7 3.4 3.4 2.0

other [M€] — — — 1.4 1.4 —

other HX [M€] 1.7 2.5 1.4 0.8 0.8 —

total CBM [M€] 1324.0 1257.1 1319.3 343.7 330.5 241.9

offsite cost [M€] 32.0 32.0 32.0 11.5 11.5 11.5

fees & contingencies [M€] 198.6 188.6 197.9 51.6 49.6 36.3

start-up [M€] 5.4 18.4 10.2 2.2 2.5 2.1

working capital [M€] 36.8 178.4 88.7 16.7 19.7 16.3

AFUDC [M€] 326.5 310.3 325.3 85.4 82.2 60.8

residual value (NPV) [M€] 2.3 -6.0 -0.7 0.2 0.0 -0.1

TCI [M€] 1925.7 1978.7 1972.7 511.3 496.0 368.7

318

Bibliography

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