· 2019-07-28 · Executive summary .................................................................................................................................4 1 Baseline

EUROPEAN COMMISSION Innovation and Networks Executive Agency

HORIZON 2020 PROGRAMME

RESEARCH and INNOVATION ACTION Call Identifier: H2020-LCE-2016-RES-CCS-RIA

Competitive Low-Carbon Energy

GA No. 727600

Flexible Dimethyl ether production from biomass gasification with sorption enhanced processes

Starting date: 1st November 2016 Duration: 48 months WP 4 – Deliverable D4.1: Preliminary process simulations

Dissemination level PU

Written By Giulio Guandalini, Alessandro Poluzzi, Matteo Romano (POLIMI)

Due date 31-10-2018 Issue date 10-11-2018

FLEDGED D4.1 – Preliminary process simulations

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Track changes

Version DATE Changes

1 10-11-2018 -

2 28-07-2019 Acknowledgement section updated

WP lead institution: POLIMI Authors’ contacts Giulio Guandalini POLIMI [email protected] Alessandro Poluzzi POLIMI [email protected] Matteo Romano POLIMI [email protected]

mailto:[email protected]

mailto:[email protected]


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Contents Executive summary ................................................................................................................................. 4

1 Baseline FLEDGED plant configuration ............................................................................................ 7

1.1 Biomass pretreatment ............................................................................................................. 7

1.2 Sorption enhanced gasification (SEG) ..................................................................................... 7

1.3 Syngas cleaning and compression ........................................................................................... 7

1.4 DME synthesis ......................................................................................................................... 8

1.5 DME purification ..................................................................................................................... 9

1.6 Heat recovery steam cycle .................................................................................................... 10

1.7 Plant integration .................................................................................................................... 11

2 ASPEN Model ................................................................................................................................. 12

2.1 Thermodynamic properties ................................................................................................... 12

2.2 Non-conventional components ............................................................................................. 12

2.2.1 Reference biomass ......................................................................................................... 13

2.3 Biomass pre-treatment ......................................................................................................... 14

2.4 Sorption enhanced gasifier (SEG) .......................................................................................... 15

2.4.1 Sensitivity analysis on SEG process parameters ............................................................ 21

2.5 Syngas purification section .................................................................................................... 23

2.6 Sorption enhanced DME synthesis (SEDMES) ....................................................................... 24

2.7 Steam cycle ............................................................................................................................ 27

3 Plant performance ......................................................................................................................... 30

3.1 Key performance indicators .................................................................................................. 30

3.2 Results ................................................................................................................................... 31

3.3 Comparison with literature ................................................................................................... 33

3.3.1 Comparison with advanced benchmark plant ............................................................... 34

4 Economic analysis .......................................................................................................................... 36

4.1 Method .................................................................................................................................. 36

4.2 Results ................................................................................................................................... 36

4.2.1 Comparison with diesel: production and avoided emission costs ................................. 36

4.3 Comparison with scientific literature .................................................................................... 37

5 Sensitivity analysis ......................................................................................................................... 39

5.1 Steam to carbon ratio ............................................................................................................ 39

5.2 DME yield .............................................................................................................................. 40

6 Conclusions .................................................................................................................................... 42

7 Acknowledgment ........................................................................................................................... 44


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Executive summary An Aspen Plus model of the reference FLEDGED plant has been developed and used to calculate the mass and energy balance of the plant. Sensitivity analyses on the S/C in the gasifier and on the DME yield in the SEDMES process have been carried out. Global biomass to DME conversion efficiency (Cold gas efficiency - CGE) of 38-45%LHV combined to biomass to electricity efficiency of 8-11% have been calculated (Table S 1). Resulting “equivalent” CGE (i.e. CGE taking into account the primary energy savings associated to electricity export) of 56-64% have been obtained. Overall results obtained are in line with those reported in the literature for other biomass to DME processes based on air-blown gasification systems (Figure S 1).

Table S 1. Energy balance and economic indexes of the reference FLEDGED plant for different S/C and DME yields.

S/C in the gasifier 1.5 0.5 1.5 1.5 DME yield in the SEDMES process 90% 90% 95% 80% SEG

Biomass input to gasifier [%] 94.86 100 94.86 94.86 Biomass input to combustor [%] 5.14 - 5.14 5.14

SEDMES DME production [kg/s] 1.49 1.58 1.57 1.33 DME thermal output [MW] 42.92 45.42 45.28 38.23

Overall CGESEG [%] 67.58 71.76 67.58 67.58 CGEglobal [%] 42.92 45.42 45.28 38.23 CGEeq [%] 58.28 64.20 58.80 56.58 Electric efficiency [%] 8.96 9.95 7.82 11.03 Net electric output, Pel [MW] 8.96 9.95 7.82 11.03 Steam turbines electric output [MW] 6.34 6.96 6.27 6.50 ICE electric output [MW] 8.48 8.87 7.40 10.56 Electric consumption of auxiliaries [MW] -5.86 -5.88 -5.86 -6.03

Economic indexes DME production cost [€/GJ] 49.6 46.1 47.5 54.8 Fixed cost [€/GJ] 21.7 20.3 20.6 24.2 Variable cost [€/GJ] 28.0 25.8 26.9 30.6


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Figure S 1. Comparison of performance indexes between the reference FLEDGED process and

selected literature, considering different technologies for biomass gasification (Air = air-fed gasfier, Oxy = oxy-fed gasifier) and DME synthesis (OT = Once-through, Rec = Products recirculation)

From the economic analysis of a 100 MWLHV biomass input plant, a DME cost of 49.6 €/GJLHV has been calculated, largely associated to the cost of biomass (unit cost of 125 €/t) and from the plant cost (216 M€), as shown in Figure S 2.

Figure S 2. Breakdown of the DME production cost of the reference FLEDGED plant (S/C=1.5, DME

yield=90%)

From the preliminary techno-economic study presented in this deliverable, the following main conclusions can be listed:

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Air/OTAir/RecOxy/OTOxy/RecFLEDGED cases

CGEeq = 40 %

60 %80 %

100 %

Capital20.83

Biomass29.91

Utilities0.95

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• The relatively low biomass to DME efficiency is largely due to the high methane content in the syngas produced in the fluidized bed gasifier. To significantly improve the DME production, a methane reforming step must be included in the plant.

• Low S/C are preferable from the energy and economic point of view, as low steam input reduced the need of heat in the gasifier. The effect of reduced S/C on the conversion of biomass must however be verified through experiments.

• High DME yield is also important to increase the overall plant efficiency and reduce the DME cost. The achievable DME yield per pass with the SEDMES process must be determined in the second part of the project through experimental activity. Gas recycle to improve overall DME yield is possible only in combination with methane reforming avoiding the build-up of inert CH4 in the loop.


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1 Baseline FLEDGED plant configuration This report describes the reference FLEDGED plant configuration and modelling approach. The process configuration and process units have been chosen based on a literature survey and discussions within the consortium. 1.1 Biomass pretreatment

Biomass pretreatment [1, 2] includes a tube bundle drier where biomass flows through the external side of a tube bundle heated by LP steam (130-140°C). Tube bundle dryers are highly efficient: drying occurs with low air flow rate, as air is required only for removing the evaporated water and not as heat carrier. Thus, sensible heat losses in the off-gas, gas flow rate to be treated and emissions are low. In this work, the air flow rate is determined to achieve 80% relative humidity in the off-gas. This is sufficient to provide enough driving force to the drying process and to avoid water condensation in the gas treatment units and ducts. The assumed outlet gas temperature has been fixed to 80°C, which is sufficiently low to avoid the risk of biomass ignition and to limit heat losses. 1.2 Sorption enhanced gasification (SEG)

Sorption enhanced gasifier process includes a dual-fluidized bed system: a gasifier/carbonator and a combustor/calciner. In the FLEDGED configuration, the first one is a bubbling fluidized bed, the second one a circulating fluidized bed. In the gasifier, both gasification reactions and CaO carbonation occur producing a syngas with a module defined as:

𝑀𝑀 =𝑁𝑁𝐻𝐻2 − 𝑁𝑁𝐶𝐶𝑂𝑂2𝑁𝑁𝐶𝐶𝑂𝑂 + 𝑁𝑁𝐶𝐶𝑂𝑂2

(1)

that has to be adjusted in order to match the requirements of downstream process. In SEG, this quantity can be varied acting on the carbon separation efficiency, which is influenced by the gasifier temperature, the sorbent flow rate and the CaO activity. Carbonated sorbent (CaCO3) and residual CaO, together with unconverted solid carbon (char) and ashes from biomass, flow to the combustor. In this reactor, char (and additional biomass if required) are burnt to provide the necessary thermal input to the system and limestone is calcined to be regenerated for the next sorption cycle. Steam is fed to the gasifier as gasifying and fluidizing agent. Air is fed to the combustor to oxidize the char from the gasifier and the additional biomass. Combustion air is preheated against combustor flue gas in a tubular heat exchanger. A minimum solid purge from the bottom bed is required to avoid alkali and ash accumulation in the system. Therefore, solids are removed from the combustor, which is energetically preferable compared to extraction from the gasifier, to limit carbon losses. Limestone makeup is needed to compensate the bed material lost in the purge. The need of inert gases in the plant is also considered for (i) sealing of the biomass feeder, (ii) solids purge extraction and (iii) filters cleaning. 1.3 Syngas cleaning and compression

Syngas from the SEG gasifier is initially cooled, recovering the heat in a bottoming steam cycle. Despite the relatively low partial pressure of CO compared to conventional gasifier without in-situ C removal, metal dusting may occur in the syngas cooler and high temperature heat is recovered through evaporator tubes. The DME synthesis unit must be fed with syngas with virtually no pollutants (e.g. dust, sulphur, chlorine, tars), so a syngas cleaning section is mandatory. In Figure 1-1, possible configurations for the syngas purification section are presented. Firstly, solid particles are removed by ceramic candle filters. Filters are assumed to operate at moderate temperature (350-400°C). The selected temperature may


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be increased depending on the observed tar composition, in case it is not sufficiently high to avoid TAR condensation on the filter surface. In case tar are removed through thermal cracking by high purity oxygen injection, filters are operated at lower temperature.

Figure 1-1 – Proposed layouts for purification section

Tar could be removed by a water scrubber. However, the expected tar removal efficiency is only about 40-80% and the treatment of the waste water is expensive and energy consuming. A better alternative commercial process is oil scrubbing, which can remove up to 99-100% of the tars [3]. Oil scrubbing can be based either on a regenerative process, such as the OLGA process [4], or on a non-regenerative process where the tar-rich oil is burned in the SEG combustor. Thermal cracking and catalytic reforming are other possible tar removal options, which would require high purity oxygen injection. After tar removal, a caustic washing step is included to remove soluble contaminants (HCl, NH3) and for bulk H2S removal. The remaining Sulphur is removed through a liquid Redox process (LO-CAT process), in which H2S is converted into elemental sulfur and water by reaction with iron oxygen carrier. This is a commercial technology capable of removing H2S down to 1 ppm, without CO2 removal. This technology is considered preferable over Rectisol, which is strongly sensitive to low tars concentration, which form strong bonds with methanol and increase sorbent regeneration consumption. Moreover, Rectisol removes also CO2 affecting the syngas module and requiring additional regeneration energy. Rectisol also needs to operate at high pressure, which requires using high cost stainless steel compressor and intercoolers. The SEDMES reactions are favored by high pressure, so a compression unit is required. A three-stage intercooled compression is assumed with heat recovery in the intercoolers for steam cycle water preheating. After bulk sulfur removal, a glycol dehydration unit is used for absorbing water from compressed syngas. After this unit, H2S and other potential contaminants (e.g. tar, alkali) are reduced to trace concentration by a ZnO bed and an activated carbon bed respectively. 1.4 DME synthesis

Syngas is fed to the DME synthesis section at high pressure and low temperature, possibly after admixture of recirculated products (mainly methanol) from the DME purification unit. The use of an additional conventional DME reactor upstream the SEDMES section is likely to be beneficial for the process. In this way, a large fraction of the syngas is converted in this conventional direct DME reactor,

Candle / Bag filters OLGA

380 °CWater

scrubber / Caustic washing

80 °C

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30 °C

Carbon steel Compressor

Dryer(Glicole /

TSA bed allumina)

Hydrogenation + ZnO guard

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Water treatment

30 °C30 °C

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Tar

Unrecovered oil to combustor

Thermal cracking

ASU

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Candle / Bag filters

900-1200°C

250-300 °C

High dust / high temperature Candle / Bag

filters

Water scrubber /

Caustic washing

100 °C 30 °C

30 °C

Guard bedActivated charcoal

Option 1

Option 2

Rectisol is strongly sensitive to pollutants like tar (strong bonds with methanol: high energy consumption for regeneration)

Tar 0.2-1.5 g/Sm3

Removal 99-100%

Tar removal 40-80%

Water scrubber /

Caustic washing

100 °C

Guard bedActivated charcoal

Air


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while the SEDMES performs only the last part of the conversion. Thanks to the conventional converter, heat duty and the amount of sorbent and catalyst in the SEDMES reactors are lower. Water generated in the conventional DME converter should be removed before feeding the SEDMES reactors. The following three options for water separation have been identified:

• Water condensation, requiring gas cooling and reheating. • Dedicated centralized PSA/TSA unit. • TSA separation within the SEDMES reactor in a dedicated entrance zone with no catalyst.

The second option appears the preferable one because of the higher efficiency compared to low temperature condensation option and because of the possibility of regenerating the sorbent independently from the SEDMES reactors. Regeneration of the SEDMES sorbent will likely need both pressure and temperature swing. Different production-regeneration cycles and sweep gas are currently being investigated. 1.5 DME purification

After heat recovery and cooling, the DME-rich flow from SEDMES reactors must be treated to separate the final DME product. DME purity requirement for automotive applications is high (99.5%w) and specific thresholds are provided for other species (i.e. hydrocarbons, carbon dioxide and monoxide, water). A purification section is therefore required, able to separate the following streams:

• DME, as final product with the quality specified by ISO standards; • Methanol, that can be recirculated to DME synthesis reactors in order to improve the yield; • Non-condensable gases that contain inerts and residual reactants (N2, CH4, H2, CO, CO2).

According to the quantity and the heating value, this gas can be burnt to produce additional steam, fed to a power generation unit (gas turbine or Internal Combustion Engine) or sent to the SEG combustor. In absence of a reforming step, off-gas recirculation to reactors would cause significant CH4 accumulation and should be avoided.

The four-steps DME purification process proposed by Clausen [5] has been selected at this stage of the project (Figure 1-2). The first step is the condensation of the products, which is performed in a flash separator operating between -37 and -50°C. Such low temperature is reached by means of a cascade chiller (CO2 + ammonia refrigeration cycles), with an overall coefficient of performance COP of 1.2. The following purification steps are performed in distillation columns, the first removing the remaining absorbed gases, the second recovering high purity DME and the third for methanol separation and recirculation. The last column may be substituted with a side product extraction of methanol from the second column. Alternatively, after SEDMED water separation from methanol can be avoided at all: the water content is very low because of the sorbent in the reactors and water can be recirculated with methanol.

Figure 1-2 – Purification layout proposed by Clausen [5]

-50 °C30 bar

OffgasOffgas

DME Methanol

Water

54 °C10 bar

85 °C2 bar

-20 °C10 bar


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1.6 Heat recovery steam cycle

In the FLEDGED plant, heat is available for recovery from different sources: • SEG combustor flue gas; • Syngas cooling before purification section; • Conventional DME synthesis reactor cooling; • SEDMES reactor cooling; • Syngas compressor intercoolers.

Other coolers are not considered (e.g. water scrubber cooling) because of low temperature or because they are part of stand-alone integrated systems (e.g. OLGA). According to the available heat, the following design choices for a single pressure level heat recovery steam cycle were taken:

• Evaporation temperature is fixed by DME reactor temperature (240°C). Evaporation temperature of 230°C (i.e. 28 bar) seems reasonable for the expected steam cycle size (about 10 MWe).

• Evaporation only in the syngas cooler and DME section is feasible because the heat available in combustor flue gas is enough for superheating the saturated steam raised in the other heat recovery sections.

• Intercoolers and DME cooling are sufficient for water pre-heating. The heat recovery steam cycle should also provide the process steam for the biomass dryer and the gasifier. Additional steam can also be produced by a cogenerative system (gas turbine or ICE) exploiting the off-gas from the DME purification section.


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1.7 Plant integration

In Figure 1-3, the configuration of the plant resulting from the previous discussion is depicted.

Figure 1-3 – FLEDGED reference plant configuration

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2 ASPEN Model In this section, the modelling assumptions and preliminary results for each section of the plant are presented. The model is developed in Aspen Plus, version 9. The preliminary results presented in this section refer to a 100 MWLHV biomass input, before drying. 2.1 Thermodynamic properties

For the calculation of the thermodynamic properties, different models are considered for the different plant sections, in order to improve the accuracy of calculations for the different phenomena:

• The general model is the RKS-BM. • For the heat recovery steam cycle, IAPWS95 model specific for water is adopted. • For the water scrubber, ELECNRTL model is applied.

2.2 Non-conventional components

Biomass is not a conventional substance with an unambiguous set of thermodynamic properties, so it is included in Aspen simulation as a non-conventional component. This kind of components allow for the definition of properties (specific heat, density, heat of combustion) without the knowledge of the detailed chemical structure of the substance. The same approach is also used to define the ash properties. In Aspen Plus a model is available for coal and it is used also for biomass and ashes with parameters changed accordingly to the specific application. The Aspen models for coal are HCOALGEN for enthalpy and DCOALIGT for density calculations. This last is not relevant for this model and will be not discussed. HCOALGEN model requires the definition of ultimate and proximate analysis, potentially also including the Sulphur analysis (organic, pyritic, sulfate). Several sub-models and correlations are available for coal properties calculation. The general ones are selected and summarized in Table 2-1.

Table 2-1 - Non-conventional components sub-models

BIOMASS ASH

Property Aspen Option Value Aspen

Option Value

Heat of combustion 6 (HCOMB)

HHV dry basis 1 (BOIEC) Standard correlation

Standard heat of formation 1 From heat of combustion 1 From heat of

combustion

Heat capacity 1 (CP1C) Fitted on correlation for biomass [6] 1 (CP1C) Standard correlation

Enthalpy basis 1 Standard state at 298.15 K and 1 bar 1

Standard state at 298.15 K and 1 bar

For biomass, the heat of combustion is given, considering the relations between LHV and HHV at different humidity:

𝐿𝐿𝐿𝐿𝑉𝑉𝑑𝑑 = 𝐿𝐿𝐿𝐿𝑉𝑉𝐴𝐴𝐴𝐴 ∙ (1 + 𝑦𝑦𝑑𝑑,𝑚𝑚) (2.1)

𝐿𝐿𝐿𝐿𝑉𝑉𝑑𝑑 = 𝐿𝐿𝐿𝐿𝑉𝑉𝑑𝑑 − �𝑦𝑦𝑑𝑑,𝑚𝑚 + 𝑦𝑦𝑑𝑑,𝐻𝐻 ∙𝑀𝑀𝑤𝑤,𝑤𝑤𝑤𝑤𝑤𝑤𝑤𝑤𝑤𝑤

𝑀𝑀𝑤𝑤,𝐻𝐻2� ∙ Δℎ𝑤𝑤𝑒𝑒𝑤𝑤 (2.2)

𝐿𝐿𝐿𝐿𝑉𝑉𝑑𝑑,𝑦𝑦𝑑𝑑,𝑚𝑚′ = 𝐿𝐿𝐿𝐿𝑉𝑉𝑑𝑑,𝑦𝑦𝑑𝑑,𝑚𝑚 + �𝑦𝑦𝑑𝑑,𝑚𝑚 − 𝑦𝑦𝑑𝑑,𝑚𝑚

′ � ∙ Δℎ𝑤𝑤𝑒𝑒𝑤𝑤 (2.3)

where 𝑦𝑦𝑑𝑑 are weight fraction dry basis and Δℎ𝑤𝑤𝑒𝑒𝑤𝑤 is 2.442 MJ/kgwater; LHV and HHV are on dry basis. A calculator is implemented to convert the 𝐿𝐿𝐿𝐿𝑉𝑉𝑑𝑑𝑤𝑤𝑦𝑦 to the input requested by the model. Properties are verified modeling the complete combustion of the biomass and comparing the calculated LHV with the input one.


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Heat capacity model follows the formulation:

𝑐𝑐𝑝𝑝,𝑑𝑑 = � �𝑎𝑎1,𝑗𝑗 + 𝑎𝑎2,𝑗𝑗𝑇𝑇 + 𝑎𝑎3,𝑗𝑗𝑇𝑇2 + 𝑎𝑎4,𝑗𝑗𝑇𝑇3� ∙ 𝑦𝑦𝑗𝑗𝑗𝑗

(2.4)

where 𝑐𝑐𝑝𝑝,𝑑𝑑 is given on dry basis, 𝑗𝑗 is each component of ultimate analysis and 𝑦𝑦𝑗𝑗 are weight fraction on dry basis. Aspen requires the parameters in appropriate units to obtain 𝑐𝑐𝑝𝑝,𝑑𝑑 in cal/gram°C with temperature in °C. The specific heat of dry wood can be estimated through the correlation:

𝑐𝑐𝑝𝑝,𝑑𝑑0 = 0.003867 ∙ 𝑇𝑇 + 0.1031 (2.5)

The heat capacity model from [6] takes into account also the additional contribution of energy absorbed by the wood-water bonds

𝑐𝑐𝑝𝑝,𝑑𝑑 =𝑐𝑐𝑝𝑝,𝑑𝑑0 + 𝑦𝑦ℎ𝑢𝑢𝑚𝑚 ∙ 𝑐𝑐𝑝𝑝,𝑤𝑤𝑤𝑤𝑤𝑤𝑤𝑤𝑤𝑤

1 + 𝑦𝑦ℎ𝑢𝑢𝑚𝑚+ 𝐴𝐴 (2.6)

where 𝑐𝑐𝑝𝑝,𝑑𝑑0 is the specific heat of dry wood, 𝑐𝑐𝑝𝑝,𝑤𝑤𝑤𝑤𝑤𝑤𝑤𝑤𝑤𝑤 is the specific heat of water (about 4.186 kJ/kgK)

and A is a correction term: 𝐴𝐴 = (0.02355 ∙ 𝑇𝑇 − 1.326 ∙ 𝑦𝑦ℎ𝑢𝑢𝑚𝑚 − 6.191) ∙ 𝑦𝑦ℎ𝑢𝑢𝑚𝑚 (2.7)

with 𝑇𝑇 in K. This correction is valid between 280 and 420 K; above this temperature (fiber saturation temperature), the simple law of mixture is valid (A = 0). Parameters 𝑎𝑎𝑖𝑖,𝑗𝑗 are regressed on this correlation, obtaining the results in Table 2-2. Error between (2.4) and (2.6) is about 2% between 10% and 50% water content.

Table 2-2 – Coefficients for biomass specific heat

𝑎𝑎1,𝑗𝑗 𝑎𝑎2,𝑗𝑗 𝑎𝑎3,𝑗𝑗 𝑎𝑎4,𝑗𝑗 Biomass 0.5712 0.00257 - - Moisture -6.9561 0.03768 - -

For ashes, the standard Aspen ashes is assumed. 2.2.1 Reference biomass

A reference woody biomass from literature has been assumed [7]. The reference values were modified in order to take into account the different moisture content with respect to the assumptions for FLEDGED input biomass. Properties are reported in Table 2-3.

Table 2-3 - Properties of as-received biomass

Reference FLEDGED model

LHV [MJ/kgAR] 15.376 9.747 HHV [MJ/kgAR] 16.924 11.574 Moisture [%w] 20.0 45.0 Proximate analysis [%w,dry]

Fixed Carbon 18.84 Volatile matter 80.0 Ash 1.16

Ultimate analysis [%w,dry] Carbon 51.19 Hydrogen 6.08 Nitrogen 0.2 Chlorine 0.05 Sulphur 0.02 Oxygen 41.3 Ash 1.16


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2.3 Biomass pre-treatment

Drying of biomass is calculated with the specific Aspen model “Dryer”. Outlet temperature and moisture content are specified. Air flow rate is adjusted in order to obtain a fixed outlet air relative humidity. The heat duty required for water evaporation is provided by the condensation of steam bled from the steam turbine. Steam flow rate is controlled by a design-spec. Additional steam can be produced exploiting the combustion of exhaust gases from SEDMES section. Assumptions and results for the case study are presented in Table 2-4 to Table 2-7.

Figure 2-1 –Biomass dryer streams and ASPEN model

Table 2-4 - Properties of as-received and dried biomass

Stream # 1 2 ASPEN stream BIO-AR BIO-DRY Mass flow [kg/s] 10.268 7.059 Temperature [°C] 25.0 80.0 LHV [MJ/kgAR] 9.739 15.264 HHV [MJ/kgAR] 11.691 16.832 Moisture [%w] 45.0 20.0

Table 2-5 - Thermodynamic properties (air and steam)

Stream # 6 7 3 4 5 ASPEN stream DR-ST DR-ST2 DRY-AIR1 DRY-AIR3 DRY-AIR4 Mass flow [kg/s] 1.185 1.185 8.840 12.049 12.049 Molar flow [kmol/s] 0.66 0.066 0.306 0.484 0.484 Temperature [°C] 299.4 153.8 25.0 80.0 80.7 Pressure [bara] 6.0 5,99 1,01 1,01 1,02 Vapor fraction [-] 1.0 0.0 1.0 1.0 1.0

Dryer1

Biomass

6 73

4

5

AirSteam

2


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Stream # 6 7 3 4 5 ASPEN stream DR-ST DR-ST2 DRY-AIR1 DRY-AIR3 DRY-AIR4 Composition [%vol]

H2O 100 100 1 37.40 37.40 N2 - - 77.30 48.88 48.88 O2 - - 20.70 13.09 13.09 CO2 - - 0.03 0.02 0.02 Ar - - 0.97 0.61 0.61

Table 2-6 - Assumptions

Assumption Component/flow Value Steam pressure DR-ST 6 bar Pressure drop HX-DRYER 0.01 bar Biomass moisture at dryer outlet DRYER 20 %w,AR Biomass temperature at dryer outlet DRYER 80°C Air relative humidity at outlet DRY-AIR4 80% Isentropic efficiency blower CMP-DRY 90%

Table 2-7 - Calculation logic

Target Manipulated variable Design spec Dryer heat balance Steam flow rate (DR-ST) DS-DRYHX Outlet air humidity Air flow rate (DRY-AIR1) DS-DRAIR

2.4 Sorption enhanced gasifier (SEG)

The SEG section is described by a 0D model, defined in agreement with the model developed by CSIC (MS06 – Lumped parameter 0D model for the gasifier). The gasifier/carbonator model is implemented in a dedicated Calculator block (C-GASS) which performs mass balances and calculates the equilibrium of the chemical reactions (i.e. WGS and carbonation) at a given temperature. The resulting yield for each specie is provided as input to a R-Yield reactor (GASS) that calculates the energy balance. The calculator requires as input:

• Composition and mass flow rate of the biomass • Composition and mass flow rate of the recirculated solids from combustor/calciner • Mass flow rate of other inlet flows (i.e. steam, additional oxygen, inerts) • Gasifier operating temperature • Flow rate of fresh limestone fed to the system (F0) in the combustor/calciner • Steam-to-carbon molar ratio at gasifier inlet

The model considers the following phenomena in the gasifier/carbonator:

• Water Gas Shift reaction (WGS)

𝐶𝐶𝐶𝐶 + 𝐿𝐿2𝐶𝐶 → 𝐶𝐶𝐶𝐶2 + 𝐿𝐿2 (2.8)

According to experimental evidence, this reaction it is not at equilibrium and the parameter 𝑝𝑝𝛿𝛿𝑊𝑊𝑊𝑊𝑊𝑊 is defined to indicate the approach to the equilibrium composition, following the approach proposed in [8].

𝑝𝑝𝛿𝛿𝑊𝑊𝑊𝑊𝑊𝑊 = 𝑙𝑙𝑙𝑙𝑙𝑙10[𝛱𝛱𝑖𝑖𝑝𝑝𝑖𝑖

𝜈𝜈𝑖𝑖

𝐾𝐾𝑝𝑝(𝑇𝑇)] (2.9)


16/44

𝐾𝐾𝑝𝑝(𝑇𝑇) = �−2.4198 + 0.0003855 𝑇𝑇 +2180.6𝑇𝑇

� (2.10)

The numerator of equation (2.9) is related with partial pressure of species in the syngas, while the denominator is the equilibrium constant as function of temperature. The value of 𝑝𝑝𝛿𝛿𝑊𝑊𝑊𝑊𝑊𝑊 from literature (see [9]) is about -0.45 at 645°C; its dependence on temperature will be fitted on experimental data. The following preliminary correlation is used at this stage:

𝑝𝑝𝛿𝛿𝑊𝑊𝑊𝑊𝑊𝑊 = � 0.01418 𝑇𝑇 − 9.9032, 𝑇𝑇 < 650°𝐶𝐶0.004708 𝑇𝑇 − 3.74374, 𝑇𝑇 ≥ 650°𝐶𝐶 (2.11)

The solution of these equations gives a relation between CO, CO2, H2 and H2O in the syngas. • Methane content in the syngas is given as a methane production per unit of dry biomass fed

to the gasifier (0.054 kgCH4/kgbiom,dry). • CxHy content is related to the methane content according to a linear function of temperature:

𝑛𝑛𝐶𝐶𝑥𝑥𝐻𝐻𝑦𝑦𝑛𝑛𝐶𝐶𝐻𝐻4

= −0.000391 𝑇𝑇 + 0.425841 (2.12)

At this stage, CxHy are considered to be only C2H4 for simplicity. • The ratio between the unconverted char flowing from the gasifier to the combustor and the

total biomass carbon content is expressed with the following relation:

𝑋𝑋𝑐𝑐ℎ𝑤𝑤𝑤𝑤 = �𝑇𝑇 − 600

840− 600(29.7− 59.4) + 59.4� ∙

1100

(2.13)

Char is assumed to be pure carbon. This parameter strongly influences the gasifier balances and its dependence on temperature, solid residence time in gasifier and other operating parameters will be included in the model once sufficient experimental data will be available. This empirical correlation has been regressed on experimental data and its extrapolation to different conditions could be unreliable, so the amount of residual char is capped to the amount of fixed carbon in the biomass.

• CO2 uptake by CaO may be limited either by the chemical equilibrium or by the maximum conversion of CaO into CaCO3 in the gasifier. The equilibrium CO2 concentration is calculated according to:

log10 𝑥𝑥𝐶𝐶𝑂𝑂2 = 7.079−8308𝑇𝑇

(2.14)

The maximum conversion of CaO into CaCO3 in the gasifier is calculated as function of the number of carbonation-calcination cycles N, according to [10]:

𝑋𝑋𝑤𝑤𝑒𝑒𝑤𝑤 = � 𝑋𝑋𝑁𝑁 ∙ 𝑟𝑟𝑁𝑁∞

𝑁𝑁=1 (2.15)

𝑋𝑋𝑁𝑁 =1

11 − 𝑋𝑋𝑤𝑤

+ 𝑁𝑁 ∙ 𝑘𝑘+ 𝑋𝑋𝑤𝑤 (2.16)

𝑟𝑟𝑁𝑁 =𝐹𝐹0 ∙ 𝐹𝐹𝐶𝐶𝑤𝑤𝑁𝑁−1

(𝐹𝐹0 + 𝐹𝐹𝐶𝐶𝑤𝑤)𝑁𝑁 (2.17)

To take into account kinetics limitation and the residence time distribution of the particles in the gasifier, a ratio between actual and maximum conversion (𝑋𝑋𝑤𝑤𝑒𝑒𝑤𝑤) of 0.75 is assumed for the sorbent. In the model, CO2 uptake is computed based on the most limiting parameter between sorbent conversion and CO2 equilibrium concentration.


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• The content of other species in the syngas are estimated assuming that all the sulfur in biomass is converted to H2S, all chlorine to HCl and 10% of the nitrogen to NH3.

The gasifier/carbonator is a bubbling fluidized bed, so the model considers an almost complete separation of solids from gas. The solids entrained by gases out of the cyclones are assumed to be 0.01% of the circulating solids. Solids recirculated are fed to the combustor/calciner that is modeled with a RGibbs reactor which performs complete combustion of char and additional biomass, as well as calcination of limestone. The reactor operates at given outlet temperature and biomass input is adjusted to close the energy balance. Because of the non-conventional nature of biomass that cannot be managed by RGibbs model, a fictitious reactor DCM-BIO is included to convert biomass into conventional species (e.g. H2, O2,…). The combustor/calciner is a circulating fluidized bed, so the solids are assumed to be entrained by the gas with a given GS (solid mass flow per cross-section unit) at raiser outlet. After solids separation by the outlet cyclone with a given efficiency, separated solids are split into the following three streams:

• a fraction, proportional to the inlet biomass flow rate, is purged in order to avoid alkali accumulation in the reactor;

• the solids flow required by the gasifier/carbonator to close the energy balance is recirculated to the gasifier;

• the remaining solids are recirculated to the CFB combustor riser. Cyclone efficiency can be specified for calcium and for ashes separately, in order to take into account the different particle size distribution. Oxygen for combustion is provided by ambient air that is preheated by flue gases. Air flow rate is adjusted to obtain a fixed oxygen excess in flue gases. The detailed assumptions and results for the case study are presented in Table 2-8 to Table 2-12.


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Gasifier - Carbonator

18

12

16

Combustor - Calciner

Dryer1

Biomass9

Air preheater

15

23 24

Air22

19

20

21

Stack

13

14

LimestoneInert solids

25

Purge10

11

8

Flue gases cooling

3

4

5

Air

2 Steam

17

26Inerts

Figure 2-2 – ASPEN model of sorption enhanced gasifier (SEG) section


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Table 2-8 - Streams properties

Stream # 2 8 9 26 10 11 ASPEN stream BIO-DRY BIO-CMB BIO-GASS INERTS STEAM-IN SYN Total mass flow [kg/s] 7.059 0.363 6.697 1.015 6.170 10.720 Temperature [°C] 80.0 80.0 80.0 25.0 170.3 720.2 Pressure [bara] 1.01 1.01 1.01 1.30 1.50 1.02 Gas/liq mass flow [kg/s] - - - 1.015 6.170 10.717 Molar flow [kmol/s] - - 0.052 0.342 0.682 Vapor fraction [-] - - - 0.14 1.00 1.00 Gas composition [%vol]

H2O - - - 86.08 100.00 57.92 H2 - - - - - 27.37 CO2 - - - - - 5.73 CO - - - - - 5.04 CH4 - - - - - 2.64 O2 - - - 2.85 - - CXHY - - - - - 0.38 Ar - - - - - - N2 - - - 11.07 - 0.90

Solid flow [kg/s] 7.059 0.363 6.697 - - 0.003 Solid composition [%w]

Biomass 100.00 100.00 100.00 - - - Ash - - - - - 74.00 CaO - - - - - 26.00 CaCO3 - - - - - - C - - - - - -

Stream # 12 13 16 17 18 25 ASPEN stream SOL-REC1 MKP-CA SOL-REC2 CMB-REC SOL-REC3 PRG-CMB Total mass flow [kg/s] 87.603 0.357 143.895 59.407 84.420 0.067 Temperature [°C] 720.2 25.0 910.0 910.0 910.0 910.0 Pressure [bara] 1.02 1.00 1.02 1.02 1.02 1.02 Gas/liq mass flow [kg/s] - - - - - - Molar flow [kmol/s] - - - - - - Vapor fraction [-] - - - - - - Gas composition [%vol]

H2O - - - - - - H2 - - - - - - CO2 - - - - - - CO - - - - - - CH4 - - - - - - O2 - - - - - - CXHY - - - - - - Ar - - - - - - N2 - - - - - -

Solid flow [kg/s] 87.603 0.357 143.895 59.407 84.420 0.067 Solid composition [%w]

Biomass - - - - - - Ash 5.38 - 5.49 5.49 5.49 5.49 CaO 88.03 - 94.51 94.51 94.51 94.51 CaCO3 5.43 100.00 - - - - C 1.15 - - - - -


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Stream # 15 19 20 21 22 23 24 ASPEN stream FLUE FLUE2 FLUE3 FLUE4 OXY-IN OXY-IN2 CMB-OXY Total mass flow [kg/s] 19.971 19.971 19.971 19.971 16.137 16.137 16.137 Temperature [°C] 910.0 300.0 120.0 134.1 25.0 44.6 271.0 Pressure [bara] 1.02 1.00 0.99 1.10 1.01 1.21 1.20 Gas/liq mass flow [kg/s] 19.755 19.755 19.755 19.755 16.137 16.137 16.137 Molar flow [kmol/s] 0.622 0.622 0.622 0.622 0.559 0.559 0.559 Vapor fraction [-] 1 1 1 1 1 1 1 Gas composition [%vol]

H2O 2.95 2.95 2.95 2.95 1.00 1.00 1.00 H2 - - - - - - - CO2 23.73 23.73 23.73 23.73 0.03 0.03 0.03 CO - - - - - - - CH4 - - - - - - - O2 3.01 3.01 3.01 3.01 20.70 20.70 20.70 CXHY - - - - - - - Ar 0.87 0.87 0.87 0.87 0.97 0.97 0.97 N2 69.44 69.44 69.44 69.44 77.30 77.30 77.30

Solid flow [kg/s] 0.216 0.216 0.216 0.216 - - - Solid composition [%w]

Biomass - - - - - - - Ash 36.97% 36.97% 36.97% 36.97% - - - CaO 63.03% 63.03% 63.03% 63.03% - - - CaCO3 - - - - - - - C - - - - - - -

Table 2-9 - Gasifier operating conditions

Value Syngas module 2.00 Syngas CO2/CO 1.137 Syngas H2/CO 5.428 Captured CO2 20.83% of inlet C Char to combustor 36.80% of inlet C Xcarb 3.34% Xcarbmax 9.04% F0/FC 0.016

Table 3-10 - Pollutants estimate

Pollutant Concentration H2S 116.40 ppmvd HCl 263.19 ppmvd NH3 42.78 ppmvd


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Assumption Component/flow Value Steam to carbon GASS 1.5 Gasifier thermal losses GASS 1.0% of total thermal input CaO conversion / Maximum CaO conversion GASS 75% Cyclones separation efficiency GASS-CYC / CMB-CYC Ca: 99.9% Ash: 99% Combustor thermal losses CMB-CALC 1.0% of total thermal input Minimum combustor temperature CMB-CALC 910 °C Economizer outlet temperature FLUE2 300°C Flue gas to stack temperature FLUE3 120°C Residual oxygen fraction in flue gases FLUE 3%vol Oxidant preheater pressure drop HX-OXYPR 0.01 bar Combustor pressure drop CMB-CALC 0.18 bar Total solid purge PRG-CMB 1% of inlet biomass Solid flux at combustor outlet (Gs) CMB-CALC 6 kg/m2s Gas superficial velocity at combustor outlet CMB-CALC 2.5 m/s Additional inert gases to gasifier INERTSref* H2O 0.85 kg/s

O2 0.0462 kg/s N2 0.1738 kg/s

Air/Flue gases blower isentropic efficiency BL-OXY1/BL-OXY2 80% Air/Flue gases blower electric efficiency BL-OXY1/BL-OXY2 94%

*Inerts proportional to the inlet biomass to the gasifier (INERTSi = INERTSref BIO-GASS/BIO-REF).

Table 2-12 - Calculation logic

Target Manipulated variable Design spec Adiabatic gasifier Solids recirculation rate Q-GASS Syngas composition (module) Gasifier temperature M-REC Combustor temperature Biomass to combustor Q-CMB Oxygen excess in combustor Oxidant flow OXY-CMB

2.4.1 Sensitivity analysis on SEG process parameters

Figure 2-3 and Figure 2-4 show two examples of sensitivity analyses that can be performed by the model. The first one calculates the influence of temperature on the syngas module, considering different heat losses from the gasifier (from 2.5% to 10% of the thermal input of biomass). Given the same temperature, additional heat losses increase the calcium recirculation to fulfill the energy balance in the gasifier, increasing in this way the carbon capture and the module. The equilibrium conditions (i.e. the carbonation reaction limited by equilibrium instead of by sorbent availability) are also plotted. At high gasifier temperature, the calculated module is independent of the heat losses, because the additional sorbent recirculated does not influence the amount of CO2 absorbed. Data from the first experiments carried out at USTUTT are plotted, showing operating conditions in the equilibrium limited region also at low temperatures, which may be caused by high heat losses from the experimental facility, requiring higher solids circulation. The second set of results (Figure 2-4) shows a similar trend when heat losses from the gasifier are constant (1%) and the ratio between actual and maximum CaO conversion (see equation (2.15)) is changed. From the figure it can be observed that higher activity lead to higher CO2 removal and higher module for a given temperature, only when CO2 uptake is limited by sorbent availability. For high gasifier temperature, where CO2 uptake is limited by the equilibrium, syngas composition is unaffected by the sorbent activity.


22/44

Figure 2-3 – Syngas module at different heat losses from gasifier.

Figure 2-4 – Syngas module estimated at different maximum calcium activity (ratio to theoretical

average stationary maximum conversion).

0

2

4

6

8

10

12

14

600 620 640 660 680 700 720 740

Syng

as m

odul

e

Gasifier temperature [°C]

10,0%7,5%5,0%2,5%EquilibriumData USTUTT

Heat losses fraction

0

2

4

6

8

10

12

14

600 620 640 660 680 700 720 740

Syng

as m

odul

e

Gasifier temperature [°C]

60%

75%

90,0%

Equilibrium

Data USTUTT

Calcium activity


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2.5 Syngas purification section

Syngas purification section is not modeled in detail as regards the OLGA and the H2S removal systems, as they are commercial modules. These units are considered as black boxes with given heat balance and electric duty.

Figure 2-5 - Syngas cleaning and compression section

Syngas compressor is assumed to be a 3-stage intercooled compressor, with intercoolers outlet temperature of 40°C. Pressure ratio per stage βstage is about 3, leading to gas temperature at each stage outlet below 200 °C.

Scrubber

Liquid redox H2S removal

Compressor

OLGA

Filter

Air/steam

Waste water

AirSulphur slurry

Waste waterRich glycol

Glycol

Activated carbon

ZnO filter

Syngas cooler

Cooler

SEG

SEDMES


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Assumption Component/flow Value Filter operating temperature SYNGAS2 340 °C OLGA inlet temperature OLGA 340 °C OLGA outlet temperature OLGA 80°C Water scrubber outlet temperature WT-SCR 40 °C Number of compressor stages CMP-SYN 3 Intercoolers outlet temperature CMP-SYN 40 °C Compressor isentropic efficiency CMP-SYN 72%

Table 2-14 - Results

Component/flow Value Compressor consumption (25 bar) CMP-SYN 4.23 MW Compressor total cooling duty (25 bar) CMP-SYN 4.23 MW Net cooling duty after heat recovery CMP-SYN 1.37 MW Compressor stages outlet temperature CMP-SYN 167-181 °C

2.6 Sorption enhanced DME synthesis (SEDMES)

DME synthesis is performed by two sequential reactors: the first one is a conventional direct DME synthesis reactor, the second one is the SEDMES reactor. The first conventional reactor is modeled with a RGibbs reactor, performing the following reactions:

𝐶𝐶𝐶𝐶2 + 𝐿𝐿2 ↔ 𝐶𝐶𝐶𝐶 +𝐿𝐿2𝐶𝐶 (2.18)

𝐶𝐶𝐶𝐶 + 2𝐿𝐿2 ↔ 𝐶𝐶𝐿𝐿3𝐶𝐶𝐿𝐿 (2.19)

2𝐶𝐶𝐿𝐿3𝐶𝐶𝐿𝐿 ↔ 𝐶𝐶𝐿𝐿3𝐶𝐶𝐶𝐶𝐿𝐿3 + 𝐿𝐿2𝐶𝐶 (2.20)

Outlet composition is the equilibrium one at the given temperature and pressure. A temperature approach to equilibrium of 100 °C is specified for methanol dehydration reaction (2.20). After the first reactor, a fraction of water is removed before SEDMES, assuming an intermediate PSA/TSA step. SEDMES reactor is modeled assuming the equilibrium composition in the gas phase and a water adsorption rate tuned to achieve a fixed DME yield of 90% of the maximum theoretical one. SEDMES products are cooled for heat recovery and DME separation. The configuration of the purification section has been presented in Section 1.5 and includes four steps: an adiabatic flash for the first non-condensable gases separation and a cascade of three distillation columns (Figure 2-6). The cooled stream to below -50 °C is separated in a flash tank and then throttled to reduce the pressure from 25 bar to 10 bar to improve the non-condensable separation that is performed in the first separation column. After these steps, a mixture of DME/methanol/water is obtained as bottom flow, which is sent to a second column through a pump that slightly increase the pressure to 15 bar in order to improve the thermal integration of reboiler/condenser. Both columns are sized to obtain a 99% DME recovery. Recirculation ratio and distillate-to-feed are adjusted to match the required specifications. Where possible, the heat required by reboilers is provided through internal heat recovery. Condensers are cooled by heat recovery and by refrigeration cycles with a COP equal to 50% of the ideal reversed Carnot cycle one. The residual methanol from the DME purification section is recirculated to the synthesis reactor for improving the conversion to DME. Off-gases from the separation section are burnt in an engine generating electricity with 40% efficiency and recovering additional 40% of heat to generate low pressure steam used for biomass drying.


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34

Water removal

36

37

38

44 4746

DME cooling39

35

Exhaust

Methanol recirculation

Water

4145

42

40

Sorption enhanced DME synthesis

Conventional DME synthesis

DME43

Figure 2-6 – ASPEN model of DME synthesis and purification sections


26/44

Table 2-15 - Streams properties

Stream # 34 35 36 37 38 39 40 ASPEN stream SYN1 SYN2 DME-R1 DME-R2 PUR0 PUR2 PUR5 Temperature [°C] 40.0 38.6 240.0 240.0 240.0 -64.3 -44.8 Pressure [bara] 25 25 25 25 25 25 10.00 Mass flow [kg/s] 3.509 3.534 3.534 3.489 2.255 2.255 1.619 Molar flow [kmol/s] 0.284 0.285 0.237 0.234 0.082 0.082 0.038 Vapor fraction [-] 1.00 1.00 1.00 1.00 1.00 0.54 0.04 Composition [%vol]

H2O - - 3.56 2.52 0.03 0.03 0.07 H2 65.73 65.58 59.43 60.07 20.47 20.47 0.23 CO2 13.68 13.64 17.21 17.39 0.97 0.97 1.72 CO 12.11 12.08 3.61 3.64 7.93 7.93 0.61 CH4 6.33 6.32 7.59 7.68 21.98 21.98 8.10 O2 - - - - - - - CXHY - - - - - - - Ar - - - - - - - N2 2.16 2.15 2.59 2.62 7.50 7.50 0.55 DME - 0.11 4.50 4.55 40.68 40.68 87.76 Methanol - 0.11 1.51 1.53 0.44 0.44 0.96

Stream # 41 42 43 44 45 46 47 ASPEN stream OFFG3 PUR65 OFFG5 DME MET-REC1 MET-REC2 MET-REC3 Temperature [°C] -33.1 45.7 22.5 61.2 84.9 85,0 86.8 Pressure [bara] 10.00 15.00 10.00 15.00 15.00 15.00 25.00 Mass flow [kg/s] 0.105 1.514 0.741 1.489 0.025 0.025 0.025 Molar flow [kmol/s] 0.005 0.033 0.049 0.032 0.001 0.001 0.001 Vapor fraction [-] 1.00 0.00 1.00 0.00 0 0 0 Composition [%vol]

H2O - 0.08 - 0.02 - - - H2 1.92 - 34.32 - - - - CO2 14.27 - 1.62 - - - - CO 5.08 - 13.29 - - - - CH4 66.96 - 36.85 - - - - O2 - - - - - - - CXHY - - - - - - - Ar - - - - - - - N2 4.51 - 12.57 - - - - DME 7.26 98.83 1.36 99.83 51.13 51.13 51.13 Methanol - 1.09 - 0.15 48.87 48.87 48.87

Table 2-16 - DME reactor assumptions

Assumption Component/flow Value Direct DME reactor temperature R-DMECON 240 °C Direct DME reactor pressure R-DMECON 25 bara Intermediate water removal efficiency SEP-RC 30% Reactor temperature R-SEDMES 240 °C Reactor pressure R-SEDMES 25 bara Total carbon conversion to DME R-SEDMES 90%


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Table 2-17 – Calculation logic

Target Manipulated variable Design spec/calculator 90% total carbon conversion to DME Water removal DS-RADS Pressure downstream VLV-CRYO Pressure in the 1st distillation column VARYP Pressure downstream PMPCOL2 Pressure in the 2nd distillation column VARYP2

2.7 Steam cycle

The preliminary heat recovery steam cycle is depicted in Figure 2-7. It is a one pressure level cycle, including a water preheating section from compressor intercoolers, a deaerator (fed by LP turbine bleeding), economizer (recovering heat from flue gases), three evaporation sections (from DME synthesis reactors, syngas cooler and SEG section flue gases) and a superheating section (from SEG flue gases). Condensate flow rate is split among the preheaters. Temperature at economizer outlet is defined to have a slightly subcooled water before the steam drum. Steam for the gasifier is bled from the turbine. The temperature-heat diagram of the steam cycle is shown in Figure 2-8.

106

7

Condenser

6648

49

Compressor intercoolers

51 52

54

65

DME reactors cooling

Syngas cooler

58

5760

6163

15

62 64

19

55

5659

50

Water makeup53

62

Flue gases from calciner

To air preheater

Figure 2-7 – ASPEN model of heat recovery steam cycle


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Table 2-18 - Streams properties (water/steam)

Stream # 48 49 50 51 51 52 52 54 ASPEN stream ST7 WMKP ST8 ST12A ST12B ST13A ST13B ST15 Temperature [°C] 32.9 25.0 27.0 27.1 27.1 109.7 109.7 155.5 Pressure [bara] 0.05 1.01 0.05 8.00 8.00 7.20 7.20 5.50 Mass flow [kg/s] 2.068 6.170 8.238 3.932 4.306 3.932 4.306 10.094 Vapor fraction [-] 0 0 0 0 0 0 0 0

Stream # 55 56 57 58 59 60 61 62 ASPEN stream ST16 ST22 ST20 ST18 ST23 ST21 ST19 ST24 Temperature [°C] 155.9 223.9 223.9 223.9 232.0 232.0 232.0 232.0 Pressure [bara] 32.00 30.00 30.00 30.00 29.00 29.00 29.00 29.00 Mass flow [kg/s] 10.094 2.683 5.272 2.139 2.683 5.272 2.139 10.094 Vapor fraction [-] 0 0 0 0 1 1 1 1

Stream # 63 10 6 7 65 66 ASPEN stream ST1 ST-GASS STDRY1 STDRY2 ST-DEA ST6 Temperature [°C] 480.0 170.5 299.4 153.8 299.4 32.9 Pressure [bara] 28.00 1.50 6.00 5.99 6.00 0.05 Mass flow [kg/s] 10.094 6.170 1.185 1.185 0.672 2.068 Vapor fraction [-] 1 1 1 0 1 0.93

Table 2-19 – Heat recovery steam cycle assumptions

Assumption Component/flow Value Temperature at turbine inlet ST1 480°C Pressure at turbine inlet ST1 28 bara Steam bleeding pressure for dearator ST-DEA 6 bara Steam bleeding pressure for dryer ST-DRY 6 bara Steam bleeding pressure for gasifier ST-GASS 1.5 bara Turbine isentropic efficiency LPT/IPT/HPT 80% Turbine mechanical/electric efficiency LPT/IPT/HPT 98% Condenser pressure COND 0.05 bara Condenser outlet conditions COND Saturated liquid Pumps hydraulic efficiency PMP1/PMP2 80% Pumps mechanical/electric efficiency PMP1/PMP2 95% Condensate extraction pump discharge pressure PMP1 8 bara Preheater pressure losses PRH1/2/3 0.8 bar Deaerator pressure DEAR 5.5 bara Feeding pump discharge pressure PMP2 32 bara ΔT subcooling at EVA inlet ECO 10 °C EVA outlet conditions EVA1/2/3 Saturated steam


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Figure 2-8 – Temperature-heat diagram of the heat recovery steam cycle

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3 Plant performance 3.1 Key performance indicators

Several performance indicators have been defined to qualify the global performance of the plant because of the presence of multiple outputs from the process. The Cold gas efficiency (CGE) is the ratio between the chemical energy of the product stream and the chemical energy input to the process (both based on LHV). Cold gas efficiency can be evaluated for the whole plant (CGE), or for any subprocess 𝑖𝑖 (e.g. for the SEG process as CGESEG, for the SEDMES as CGESEDMES).

𝐶𝐶𝐶𝐶𝐶𝐶𝑖𝑖 =�̇�𝑚𝑜𝑜𝑢𝑢𝑤𝑤,𝑖𝑖 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝑜𝑜𝑢𝑢𝑤𝑤,𝑖𝑖

�̇�𝑚𝑖𝑖𝑖𝑖,𝑖𝑖 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝑖𝑖𝑖𝑖,𝑖𝑖 (3.1)

CGEu is introduced in order to account only for the useful species for DME production. Therefore, as regards the SEG section, only the hydrogen and the carbon monoxide are considered in the computation. For what concerns the DME synthesis, the CGEu

SEDMES considers the useful feed (i.e. hydrogen and carbon monoxide) and the useful product (i.e. DME). The same happens for the DME purification section, where CGEu

DME_pur considers the DME entering/exiting the distillation section as the useful feed/products.

𝐶𝐶𝐶𝐶𝐶𝐶𝑆𝑆𝑆𝑆𝑆𝑆𝑢𝑢 =�̇�𝑚𝐻𝐻2 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝐻𝐻2 + �̇�𝑚𝐶𝐶𝑂𝑂 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝐶𝐶𝑂𝑂

�̇�𝑚𝑏𝑏𝑖𝑖𝑜𝑜𝑚𝑚𝑤𝑤𝑏𝑏𝑏𝑏 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝑏𝑏𝑖𝑖𝑜𝑜𝑚𝑚𝑤𝑤𝑏𝑏𝑏𝑏

(3.2)

𝐶𝐶𝐶𝐶𝐶𝐶𝑆𝑆𝑆𝑆𝑆𝑆𝑆𝑆𝑆𝑆𝑆𝑆𝑢𝑢 =�̇�𝑚𝑆𝑆𝑆𝑆𝑆𝑆 𝑝𝑝𝑤𝑤𝑜𝑜𝑑𝑑 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝑆𝑆𝑆𝑆𝑆𝑆 𝑝𝑝𝑤𝑤𝑜𝑜𝑑𝑑

�̇�𝑚𝐻𝐻2 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝐻𝐻2 + �̇�𝑚𝐶𝐶𝑂𝑂 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝐶𝐶𝑂𝑂

(3.3)

𝐶𝐶𝐶𝐶𝐶𝐶𝑆𝑆𝑆𝑆𝑆𝑆_𝑝𝑝𝑢𝑢𝑤𝑤𝑢𝑢 =

�̇�𝑚𝑆𝑆𝑆𝑆𝑆𝑆 𝑜𝑜𝑢𝑢𝑤𝑤 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝑆𝑆𝑆𝑆𝑆𝑆 𝑜𝑜𝑢𝑢𝑤𝑤 �̇�𝑚𝑆𝑆𝑆𝑆𝑆𝑆 𝑝𝑝𝑤𝑤𝑜𝑜𝑑𝑑 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝑆𝑆𝑆𝑆𝑆𝑆 𝑝𝑝𝑤𝑤𝑜𝑜𝑑𝑑

(3.4)

The Electric efficiency (EE) indicates the conversion of the biomass chemical energy into electricity. The net electric output (Pel) accounts for the electricity produced by the steam turbines and by the internal combustion engine and for the electric consumption of the auxiliaries.

𝐶𝐶𝐶𝐶 =𝑃𝑃𝑤𝑤𝑙𝑙

�̇�𝑚𝑏𝑏𝑖𝑖𝑜𝑜𝑚𝑚𝑤𝑤𝑏𝑏𝑏𝑏 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝑏𝑏𝑖𝑖𝑜𝑜𝑚𝑚𝑤𝑤𝑏𝑏𝑏𝑏

(3.5)

CGEeq accounts for the biomass saving associated with the electricity production of the plant. A steam cycle with 34% of electric efficiency (ηel,ref) is assumed as a reference, considering a medium-size biomass-fed subcritical steam power plant.

𝐶𝐶𝐶𝐶𝐶𝐶𝑤𝑤𝑒𝑒 =�̇�𝑚𝑆𝑆𝑆𝑆𝑆𝑆 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝑆𝑆𝑆𝑆𝑆𝑆

�̇�𝑚𝑏𝑏𝑖𝑖𝑜𝑜𝑚𝑚𝑤𝑤𝑏𝑏𝑏𝑏 ∙ 𝐿𝐿𝐿𝐿𝑉𝑉𝑏𝑏𝑖𝑖𝑜𝑜𝑚𝑚𝑤𝑤𝑏𝑏𝑏𝑏 −𝑃𝑃𝑤𝑤𝑙𝑙

𝜂𝜂𝑤𝑤𝑙𝑙,𝑤𝑤𝑤𝑤𝑟𝑟

(3.6)


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C-yield DME defines the yield of the DME exiting the fuel synthesis section with respect to the input carbon monoxide and dioxide, considering the molar flows.

𝐶𝐶 − 𝑦𝑦𝑖𝑖𝑦𝑦𝑙𝑙𝑦𝑦 𝐷𝐷𝑀𝑀𝐶𝐶 =2 ∙ 𝐹𝐹𝑆𝑆𝑆𝑆𝑆𝑆 𝑜𝑜𝑢𝑢𝑤𝑤 − 2 ∙ 𝐹𝐹𝑆𝑆𝑆𝑆𝑆𝑆 𝑖𝑖𝑖𝑖

�FCO2 + 𝐹𝐹𝐶𝐶𝑂𝑂�𝑖𝑖𝑖𝑖 (3.7)

3.2 Results

Figure 3-1 shows the energy balance of the SEG section. The largest part of the input energy is associated to the chemical energy of biomass. Minor contributions derive from the sensible heat of the other streams fed to the SEG unit. About 67% of the energy output is represented by the chemical energy of the raw syngas exiting the process. Anyway, only the 50% is useful chemical energy (i.e. LHV of CO and H2 species), which can be converted in the downstream synthesis process. More than 16% of the SEG output is associated to the heating value of methane and higher hydrocarbons which cannot be exploited for DME production. Significant contributions to the outlet energy flow are related to the sensible heat of the syngas exiting the gasifier and of the combustor flue gases.

Figure 3-1 – SEG process energy balance. Inner circle: input energy. Outer circle: output energy.

The performance indicators which have been described in the previous section together with other key quantities are reported in Table 3-1 for the different plant sections. The gasifier temperature is adjusted to obtain a module of 2 at gasifier outlet for the downstream fuel synthesis. In the assessed case, about 5% of the biomass is fed to the combustor in order to keep it adiabatic. The gross electricity production, given by the steam turbine and by the internal combustion engine, exceeds the electric process consumption. Therefore, the excess electricity is exported to the electric grid.

LHV biomass97.43%

Sensible heat biomass0.72%

Sensible heat steam1.66%

Sensible heat inerts0.04%

Sensible heat oxydant0.15%

LHV CO and H250.26%

LHV other gases16.38%

Sensible heat syngas15.76%

Stack losses2.17%

Heat purge0.05%

Thermal losses1.83%

Flue gases cooler12.96%

First calcination0.58%

Combustor cooling0.00%


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Table 3-1 - Plant performance

Performance indexes Dryer

Dryer heat input [MW] 9.46 SEG

Gasifier temperature [°C] 720.2 FCa/Fc 6.23 CaO conversion in the gasifier [%] 3.34 Thermal energy extracted from the combustor [MW] - Biomass input to gasifier [%] 94.86 Biomass input to combustor [%] 5.14 CGESEG [%] 67.58 CGEuSEG [H2, CO] [%] 50.94

SEDMES DME production [kg/s] 1.49 DME output [MWLHV] 42.92 CGESEDMES [%] 93.42 CGEuSEDMES [H2, CO] [%] 80.67 CGEDME_pur [%] 66.29 CGEuDME_pur [DME] [%] 97.03

Steam cycle Net electric efficiency [%] 20.69

Overall indicators CGEglobal [%] 42.92 EE [%] 8.96 CGEeq [%] 58.28 Net electric output, Pel [MW] 8.96 Steam turbines electric output [MW] 6.34 ICE electric output [MW] 8.48 Electric consumption [MW] 5.86

Dryer compression consumption 0.01 SEG air fan 0.34 SEG flue gas fan 0.30 Pump water scrubber 0.01 OLGA 0.09 LIQ-REDOX 0.03 Syngas compressor 4.23 SEDMES consumption 0.81 Steam cycle pumps 0.04

Table 3-2 shows the evolution of the CGE along the plant. CGEtot represents the cumulative CGE, calculated as the product between the current CGEi (or CGEu

i) and the one of the preceding unit.

Table 3-2 – Evolution of CGE along the plant units.

CGE [%]

Drying SEG Cleaning DME synthesis DME purification

CGEi 107.7% 67.6% 95.2% 93.4% 66.3% CGEtot - 72.8% 69.3% 64.8% 42.9% CGEui 107.7% 50.9% 100.0% 80.7% 97.0% CGEutot - 54.9% 54.9% 44.3% 42.9%


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3.3 Comparison with literature

In Table 3-2, selected performance indicators are used to compare the simulated plant with values presented in literature for different DME production processes. In Figure 3-3, the performance indicators of these plants are plotted on a CGE-EE chart, reporting the iso-CGEeq lines. The calculated reference FLEDGED plant shows CGE and CGEeq in the upper range with respect to other once-through, air-fed gasification systems.

Table 3-3 – Performance indexes comparison from literature, considering different technologies for biomass gasification (CFB = Circulating Fluidized Bed; EF = Entrained Flow; SR = Steam reformer;

ATR = Auto-thermal reformer, TR = Tar cracker/Methane reformer) and DME synthesis (Direct/Indirect, OT = Once-through, Rec = Products recirculation).

Author CGEgasif CGEplant ηel CGEeq, pl Gasifier DME synthesis

This work 72.80 42.92 8.96 58.28 SEG, air ind, OLGA, liq Redox SEDMES

Haro (SR-03) [11] 80.75 34.36 4.80 40.01 i-CFB, air ind, SR, OLGA Direct, OT

Haro (ATR-03) [11] 80.75 39.14 1.61 39.40 i-CFB, air ind, ATR Direct, OT

Haro (TR-06) [11] 80.75 33.66 8.42 44.74 i-CFB, air ind, TR, MEA Direct, Rec

Tock (DME) [12] - 48.00 <0 - CFB, air ind, SR, MEA Direct, Rec

Larson (D-OT) [13] 79.80 24.30 30.20 > 100 no dry, CFB, oxy,

TR, Rectisol Direct, OT

Larson (D-RC) [13] 79.80 52.30 8.80 70.56 no dry, CFB, oxy, TR, Rectisol Direct, Rec

Tuna [14] - 53.00 3.70 58.13 BFB, oxy, TR 2 step, Rec Wang [15] 73.90 37.31 - - CFB, air, biogas rec Direct, OT Clausen [16] 83.00 60.00 3.00 65.81 EF, oxy Rec Clausen (sm-OT) [17] 93.00 46.00 6.00 55.86 Fixed 2-stage, air OT

Clausen (sm-RC) [17] 93.00 58.00 0 58.00 Fixed 2-stage, air Rec

Clausen (lg-OT) [17] 81.00 49.00 16.00 92.56 EF, oxy, Rectisol OT

Clausen (lg-RC) [17] 81.00 66.00 5.00 77.38 EF, oxy, Rectisol Rec

Arvidsson [18] 83.00 57.00 -3.00 - Torref, EF, oxy, Rectisol Direct, Rec

Van Rens [19] 76.89 55.60 -6.19 67.98 CFB, oxy, ATR, WGS, Rectisol 2 step, Rec


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Figure 3-2 – Performance indexes comparison from literature, considering different technologies

for biomass gasification (Air = air-fed gasfier, Oxy = oxy-fed gasifier) and DME synthesis (OT = Once-through, Rec = Products recirculation).

3.3.1 Comparison with advanced benchmark plant

In addition to the previous general comparison, the plant performance indicators are compared with the 300 MWth base case design reported in [20], which represents the performance of the VTT’s Ultra-Clean Gas process. Detailed data for this system are available, allowing a more detailed comparison of FLEDGED process with this case. The plant is based on circulating fluidized-bed O2-blown gasifier operating at 5 bar, followed by hot-filtration at 550°C and catalytic reforming of hydrocarbons and tars. Then, the reformed syngas is fed to a sour shift reactor, followed by a water scrubber. Syngas is then compressed to 20 bar prior to an acid gas removal step based on Rectisol. The DME synthesis island consists in three to four intercooled reactors for large scale plants and a single reactor for small scale plant, with unconverted gases recycle loop. DME reactors are operated at 60 bar, and follow a second compression step. DME recovery and purification is performed in a distillation section. The whole process is integrated with auxiliary equipment which includes the biomass dryer, the air separation unit, an auxiliary boiler and a steam cycle. A comparison of the main performance indicators with the DME plant discussed in this report is given in Table 3-3.

Table 3-4 – Performance comparison between FLEDGED and VTT plants.

Performance FLEDGED plant VTT plant CGE: drying [%] 107.7 111.7 CGE: gasification + syngas conditioning [%] 64.4 69.6 CGE: DME synthesis [%] 61.9 76.8 CGE: global [%] 42.9 59.7 Methane slip [% MW/MWbio] 14.4 1.3 EE [%] 9.0 2.5 CGEeq [%] 58.3 64.3

The VTT case design exhibits higher CGE for each section of the plant. The high concentration of methane is the main reason for the low CGE of the plant, as also shown by the difference in the methane slip. This is due to the auto-thermal reforming section in the VTT plant, sustained by injection of oxygen form the ASU. On the other hand, in the VTT case design, the electric efficiency is much

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lower than the FLEDGED plant one. As a result, CGEeq 6% points lower is obtained with the reference FEDGED plant compared to the VTT plant.


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4 Economic analysis 4.1 Method

The economic analysis of the simulated DME plant estimates the DME production cost that results from the evaluation of the total investment cost, the maintenance cost and the utility price. Most of the Capex estimation derives from [21]. For the conventional DME reactor, the cooling plant and the DME purification section, data from [5] are used for Capex estimation. A total capital investment of 216.1 M€ has been estimated. The fixed operating and maintenance cost are assumed 4% of the capital investment. For the variable costs, the following reference costs are assumed: 125 €/ton for the biomass, 2.5 €/m3 for the demi water, 0.2 €/m3 for the cooling water make-up and 65 €/ton for the limestone. The electricity price amounts to 50 €/MWh, as reported in [20]. The plant capacity factor is set to 90%, which corresponds to 7884 annual operating hours.

4.2 Results

The calculated DME production cost with the reference assumptions is 49.6 €/GJ. Figure 4-1 illustrates the breakdown of the DME production cost. The main contribution for the DME production cost is associated to the biomass cost, followed by the capital investment cost. The excess electricity produced by the steam cycle sold to the power grid contributes lowering the overall cost. The other utilities and the operating and maintenance cost have a minor contribution.

Figure 4-1 - Different contributions to the DME production cost

4.2.1 Comparison with diesel: production and avoided emission costs

The estimated DME production cost is compared with the production cost of diesel which amounts to 17.8 €/GJ [22]. When taxes are included, the diesel cost increases to 38.4 €/GJ. Both the quantities result from a weighted average of market values among the 28 countries of the European Union. Figure 4-2 shows DME iso-cost lines on a total capital investment-biomass cost chart. The DME production cost of the FLEDGED reference plant (red point) is above the lines corresponding to the diesel cost. The chart shows for example that with the estimated capital investment costs, a biomass price of about 80 €/t is needed to make the DME produced with the FLEDGED process competitive with the conventional diesel (after tax).

Capital20.83

Biomass29.91

Utilities0.95

Power-2.90

O&M0.83

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Figure 4-2 – DME production cost equal to diesel production cost breakeven lines (blue lines). DME production cost of the FLEDGED reference plant (red point)

As previously shown, the cost of the utilities is the most important factor which contributes to the DME cost. Among the different utilities, biomass gives the largest contribution. Figure 4-3 shows the DME production cost for different biomass prices. Furthermore, since the total investment cost is highly uncertain, the DME production cost is plotted showing the effect of a capital cost variation by ±30%.

Figure 4-3 – DME production cost as function of biomass price with a capital cost variation by ±30%

4.3 Comparison with scientific literature

In this section, a comparison with the economic analysis of the VTT plant reported in [20] and described in 3.3.1 is performed. The DME production cost estimated in [20] is 18.4 €/GJ, which is significantly lower than the one computed in this work. The main reason of such a difference is due to the lower

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biomass price assumed in [20], which is 40 €/t. If the same biomass cost is assumed, DME production cost of the FLEDGED plant reduces to 29.3 €/MJ. Table 4-2 shows a comparison between the specific capital cost of the two plants, referred to the kWLHV of biomass input. In all the considered sections of the process, the VTT plant exhibits lower investment cost than the FLEDGED case. It is worth noting that the VTT plant is 3 times bigger, therefore a significant scale effect is expected.

Table 4-1 – Capital cost comparison between FLEDGED and VTT plant

Total capital investment [€/kWLHV,biom] FLEDGED plant VTT plant Auxiliary equipment + gasification + syngas purification* 1163 857 Power island 146 82 DME synthesis 664 232 Heat rejection system 37 - Waste water treatment 151 - Total 2161 1171

*The auxiliary equipment includes the biomass pretreatment and the air separation unit (only VTT case). The syngas purification section includes filters, reformer and WGS reactor (only VTT case), OLGA (only FLEGDGED case), scrubber, syngas compression, acid gas removal (liquid Redox process for FLEGDGED, Rectisol for VTT).


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5 Sensitivity analysis 5.1 Steam to carbon ratio

A sensitivity analysis is performed by decreasing the steam to carbon ratio at gasifier inlet to 0.5. The energetic and economic performance indexes are shown in Table 5-1. In this case, because of the lower sensible heat needed for steam heating, no additional biomass must be fed to the combustor, which must be cooled as the unconverted carbon flowing from the gasifier to the combustor exceeds the thermal need of the process. The lower S/C results in an increase of the CGE of the SEG section. Furthermore, the steam cycle produces more power and exhibits a higher electric efficiency, since a lower steam flow rate is bled from the turbine to be fed to the gasifier. This is beneficial for both EE and CGEeq. From the economic perspective, the reduction of S/C leads lower fixed and variable costs, resulting in a reduction of the DME production cost from 49.6 to 46.1 €/GJ.


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Table 5-1 – Sensitivity analysis on steam to carbon ratio

Performance and economic indexes S/C=1.5 S/C=0.5 Dryer

Dryer heat input [MW] 9.46 9.46 SEG

Gasifier temperature [°C] 720.2 721.18 FCa/Fc 6.23 4.57 CaO conversion in the gasifier [%] 3.34 4.54 Thermal energy extracted from the combustor [MW] - 1 Biomass input to gasifier [%] 94.86 100 Biomass input to combustor [%] 5.14 - CGESEG [%] 67.58 71.76 CGEuSEG [H2, CO] [%] 50.94 54.23

SEDMES DME production [kg/s] 1.49 1.58 DME output [MWLHV] 42.92 45.42 CGESEDMES [%] 93.42 92.68 CGEuSEDMES [H2, CO] [%] 80.67 80.20 CGEDME_pur [%] 66.29 66.56 CGEuDME_pur [DME] [%] 97.03 97.03

Steam cycle Net electric efficiency [%] 20.69 25.60

Overall CGEglobal [%] 42.92 45.42 EE [%] 8.96 9.95 CGEeq [%] 58.28 64.20 Net electric output, Pel [MW] 8.96 9.95 Steam turbines electric output [MW] 6.34 6.96 ICE electric output [MW] 8.48 8.87 Electric consumption [MW] 5.86 5.88

Dryer compression consumption 0.01 0.01 SEG air fan 0.34 0.31 SEG flue gas fan 0.30 0.27 Pump water scrubber 0.01 0.01 OLGA 0.09 0.09 LIQ-REDOX 0.03 0.03 Syngas compressor 4.23 4.28 SEDMES consumption 0.81 0.85 Steam cycle pumps 0.04 0.04

Economic indexes DME production cost [€/GJ] 49.6 46.1 Fixed cost [€/GJ] 21.7 20.3 Variable cost [€/GJ] 28.0 25.8

5.2 DME yield

Another sensitivity analysis is performed by varying the DME yield in the SEDMES process with respect to the reference value (see Table 5-2). When the yield is increased to 95% of the maximum theoretical one, more DME fuel is produced (higher CGEu

SEDMES). A higher DME production causes a lower off-gas flow which is burnt inside the ICE, decreasing its power production and, therefore, the EE of the plant. The opposite happens when the DME yield is decreased to 80%. From the economic point of view,


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when the DME yield is increased, the DME production cost slightly decreases, mainly because of a higher fuel production given the same investment.

Table 5-2 - Sensitivity analysis on C-yield DME

Performance indexes YDME=90% YDME=95% YDME=80% Dryer

Dryer heat input [MW] 9.46 9.46 9.46 SEG

Gasifier temperature [°C] 720.2 720.2 720.2 FCa/Fc 6.23 6.23 6.23 CaO conversion in the gasifier [%] 3.34 3.34 3.34 Thermal energy extracted from the combustor [MW] - - - Biomass input to gasifier [%] 94.86 94.86 94.86 Biomass input to combustor [%] 5.14 5.14 5.14 CGESEG [%] 67.58 67.58 67.58 CGEuSEG [H2, CO] [%] 50.94 50.94 50.94

SEDMES DME production [kg/s] 1.49 1.57 1.33 DME output [MWLHV] 42.92 45.28 38.23 CGESEDMES [%] 93.42 92.74 94.80 CGEuSEDMES [H2, CO] [%] 80.67 85.15 71.71 CGEDME_pur [%] 66.29 70.44 58.18 CGEuDME_pur [DME] [%] 97.03 97.03 97.03

Steam cycle Net electric efficiency [%] 20.69 20.22 21.67

Overall CGEglobal [%] 42.92 45.28 38.23 EE [%] 8.96 7.82 11.03 CGEeq [%] 58.28 58.80 56.58 Net electric output, Pel [MW] 8.96 7.82 11.03 Steam turbines electric output [MW] 6.34 6.27 6.50 ICE electric output [MW] 8.48 7.40 10.56 Electric consumption [MW] 5.86 5.86 6.03

Dryer compression consumption 0.01 0.01 0.01 SEG air fan 0.34 0.34 0.34 SEG flue gas fan 0.30 0.30 0.30 Pump water scrubber 0.01 0.01 0.01 OLGA 0.09 0.09 0.09 LIQ-REDOX 0.03 0.03 0.03 Syngas compressor 4.23 4.23 4.23 SEDMES consumption 0.81 0.80 0.98 Steam cycle pumps 0.04 0.05 0.04

Economic indexes DME production cost [€/GJ] 49.6 47.5 54.8 Fixed cost [€/GJ] 21.7 20.6 24.2 Variable cost [€/GJ] 28.0 26.9 30.6


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6 Conclusions An Aspen Plus model of the reference FLEDGED plant has been developed and used to calculate the mass and energy balance of the plant. Sensitivity analyses on the S/C in the gasifier and on the DME yield in the SEDMES process have been carried out. Biomass to DME conversion efficiency of 38-45%LHV combined to biomass to electricity efficiency of 8-11% have been calculated. Overall results obtained are in line with those reported in the literature for other biomass to DME processes based on air-blown gasification systems. From the economic analysis of a 100 MWLHV biomass input plant, a DME cost of 49.6 €/GJLHV has been calculated, largely associated to the cost of biomass (unit cost of 125 €/t) and from the plant cost (216 M€). The following main conclusions from this study can be listed:

• The relatively low biomass to DME efficiency is largely due to the high methane content in the syngas produced in the fluidized bed gasifier. To significantly improve the DME production, a methane reforming step must be included in the plant.

• Low S/C are preferable from the energy and economic point of view, as low steam input reduced the need of heat in the gasifier. The effect of reduced S/C on the conversion of biomass must however be verified through experiments.

• High DME yield is also important to increase the overall plant efficiency and reduce the DME cost. The achievable DME yield per pass with the SEDMES process must be determined in the second part of the project through experimental activity. Gas recycle to improve overall DME yield is possible only in combination with methane reforming avoiding the build-up of inert CH4 in the loop.

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[22] “Weekly Oil Bulletin,” 2018. .


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7 Acknowledgment This project has received funding from the European Union's Horizon 2020 research and innovation programme under grant agreement No 727600 http://ec.europa.eu

http://ec.europa.eu/

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· 2019-07-28 · Executive summary .................................................................................................................................4 1 Baseline