Techno-Economic Evaluation of a Small-Scale Power Generation Unit

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Kinetics, Catalysis, and Reaction Engineering

Techno-Economic Evaluation of a small-scale power generation unit based on a Chemical Looping Combustion Process in Fixed Bed Reactor network Giuseppe Diglio, Piero Bareschino, Erasmo Mancusi, and Francesco Pepe Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b02378 • Publication Date (Web): 25 Jul 2018 Downloaded from http://pubs.acs.org on August 4, 2018

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Industrial & Engineering Chemistry Research

Techno-Economic Evaluation of a small-scale power generation unit based on a Chemical Looping Combustion Process in Fixed Bed Reactor network

Giuseppe Diglio, Piero Bareschino, Erasmo Mancusi*, Francesco Pepe

Dipartimento di Ingegneria, Università degli Studi del Sannio, Piazza Roma 21, 82100 Benevento, Italy

*Corresponding Author: E. Mancusi ([email protected])

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Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Abstract This paper reports the techno-economic analysis of a small-scale power generation unit integrating a Chemical Looping Combustion (CLC) with inherent CO2 capture section and a power one. In the CLC section four adiabatic fixed beds are operated using Cu as oxygen carrier and methane as fuel to continuously produce two hot gas streams at approximately constant mass flow rate and temperature of about 1045 K and 1175 K, respectively. While the former is used to warmup the inlet air, the energetic content of the latter is converted into electricity through a gas turbine system based on a combined cycle. By means of a 1D numerical model, it was assessed that the net power generation of the proposed power unit is about 0.50 MW, with a global energy efficiency (51%) higher than that of alternative CCS technologies. The proposed system is characterised by better economic performance than alternative CCS-based power plants, reaching a levelised cost of energy and cost of CO2 avoided of about 54 €·MWh-1 and 31 €·tonCO2-1, respectively. By means of a sensitivity analysis it was assessed that the economic performances of the proposed system were primarily affected by the specific cost of fuel, while its economic feasibility mainly relies on the lifetime of a high-temperature valve. Under the considered conditions, a pay-back period of around 2.8 years and a cumulative profit in 25 years about 5 times greater than the total capital requirement were evaluated.

Keywords: Chemical Looping Combustion; Fixed Bed; Copper; Power Plant; Cost of Energy


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Nomenclature AC

Cost of CO2 avoided, €·tonCO2-1

ap

Characteristic length of solid particle (volume/surface), m

Ar

Cross-section of the reactor, m2

B

Stoichiometric coefficient

C

Gas specie concentration, molˑm-3

CAF

Cash Flow, €

CDCF

Cumulative Discounted Cash Flow, €

CF

Capacity factor

CFlowPP,CLC Cash flow of CLC power plant, € Cp

Heat capacity, J∙kg-1∙K-1

CTS

CO2 transport and storage cost, €·tonCO2-1

Dax

Effective axial dispersion, m2∙s-1

De,i

Effective diffusion coefficient of specie i, m2∙s-1

dp

Particle diameter, m

DR

Discount Rate, %

dr

Reactor diameter, m

E

Activation energy, Jˑmol-1

e

Efficiency

eCO2

Specific CO2 emission, tonCO2∙MWh-1

FCF

Fixed Charge Factor

FOM

Fixed operating and maintenance cost, €

GS

Mass flux of the gas phase, kg∙m-2∙s-1

hm

Gas-solid mass transfer coefficient, mˑs-1

K

Chemical reaction rate constant, molˑm-3ˑs-1

K0

Pre-exponential factor of chemical reaction rate constant, mol1-mˑm3m-2ˑs-1

km

Gas-solid mass transfer coefficient, W∙m-2∙K-1

L

Reactor length, m

LCOE

Levelised Cost of Energy, €·MWh-1

LHVCH4

Low heating value of methane, J∙kg-1

lt

Life time, y

m

Reaction order



M

Molecular weight, kgˑmol-1

Nmin

Minimum number of reactors

Nr

Number of reactors

NWP

Weisz-Prater number

P

Pressure, Pa

Pel

Electric power, W

PIR

Project Interest Rate, %

Pr

Prandtl number

Q

Volumetric flow rate, m3∙s-1

r

Global reaction rate, mol∙m-3∙s-1

R

Ideal gas constant, J∙mol-1∙K-1

Re

Reynolds number

s

Thickness of reacting solid, m

Sc

Schmidt number

SFC

Specific Fuel Cost, €·MWh-1

t

Time, s

T

Temperature, K

T0

Fixed bed initial temperature, K

TCR

Total Capital Requirement, €

ug

Gas superficial velocity, m∙s-1

VCTS

Variable operating and maintenance CO2 transport and storage cost, €·MWh-1

VOM

Variable operating and maintenance cost, €

VS

Variable operating and maintenance solid material cost, €·MWh-1

VV

Variable operating and maintenance high-temperature valve cost, €·MWh-1

wact0

Mass fraction of the active phase on the carrier

wCO2,out

CO2 outlet mass fraction

wr,g

Inlet mass fraction of reacting gas

X

Solid conversion

y

Gas specie molar fraction

z

Axial spatial variable, m

Greek letters

Π

Period, s

∆H

Reaction enthalpy, J∙mol-1


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∆T

Maximum temperature increase, K

α

Partition ratio

β

Fraction of GS,RS supplied by the thermal buffer

ε

Void fraction

γ

Amortisation years

η

Reaction effectiveness

λ

Thermal conductivity, W∙m-1∙K-1

λax

Effective heat dispersion coefficient, W∙m-1∙K-1

λe0

Static contribution to effective heat dispersion (0.01 W∙m-1∙K-1)

µ

Viscosity, Pa∙s

ν

Percentage of the whole CLC period, %

ρ

Density, kgˑm-3

ρm,k

Solid molar density, molˑm-3

τ

Reaction time for complete solid conversion, s

Subscripts act

Active phase

com

Compressor

disp

Displacement

g

Gas

i

Gas specie index (CH4, H2O, CO2, O2, N2)

in

Inlet

j

Reaction index (R1, R2)

k

Solid specie index (Cu, CuO)

max

Maximum

net

Net

obs

Observed

p

Particle

r,g

Reacting gas

s

Solid material

v

Valve

Acronyms CCS

Carbon Capture and Storage 5 Environment ACS Paragon Plus


CLC

Chemical Looping Combustion

DFB

Dual Fluidised Bed

HROS

Heat Removal after Oxidation Stage

HRRS

Heat Removal after Reduction Stage

IGCC

Integrated Gasification Combined Cycle

NG

Natural Gas-fired power plant

OC

Oxygen Carrier

OS

Oxidation Stage

PP

Power Plant

PP,CLC

CLC Power Plant

PS

Purge Stage

REF

REFerence system

RS

Reduction Stage

SC-PC

SuperCritical-Pulverized Coal


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1. Introduction In the recent Paris Conference of the Parties an important agreement was reached to limit CO2 concentration in lower atmosphere at 450 ppm by 2100, with the aim to keep “well below 2°C” the increase in average world temperature1. This policy measure led to more and more interest in decarbonisation of energy sector, which represents by far the main contributor to atmospheric CO2 emissions2. Renewable energies are the main alternative to sustainably satisfy worldwide primary energy demand, but not in the short and middle term3, thus fossil fuels will be still used at least over the next few decades. In this context, Carbon Capture and Storage (CCS) technologies represents the best option to address climate challenges. The main drawback of CCS is the energy penalty associated with CO2 capture, thus great efforts were carried out during the last years to develop new high-efficiency CCS technologies, Chemical Looping Combustion (CLC) appearing one of the most promising ones4. CLC is a cyclic unmixed combustion process based on an Oxygen Carrier (OC), usually a metal oxide impregnated on an inert support, able to be alternatively oxidized by air and reduced by fuel. Using this technology the direct contact between atmospheric oxygen and fuel is avoided, with the important feature that a N2-free stream of concentrated CO2 ready for storage can be easily obtained. CLC processes are typically characterized on the basis of both reactor configuration and oxygen carrier. With respect to the former, Dual Fluidised Bed (DFB) and fixed bed are the main arrangements; while in the first OC is continuously re-circulated between an air reactor, where carrier oxidation occurs, and a fuel reactor, where fuel oxidation takes place at the expenses of carrier reduction, in the second the solid is fixed and cyclically exposed to oxidizing and reducing conditions by switching feed gas streams5. Although DFB is the most used arrangement, interest towards fixed bed is recently increasing6. With respect to the latter, a number of OCs were proposed4, with copper-based one being a suitable candidate due its good oxygen transport capacity. CLC in fixed bed using Cu-based OC was extensively studied from both experimental and numerical point of view. Kooiman et al.7 experimentally investigated a CLC process consisting of two pressurized fixed beds in series, with Cu-based OC in the first one and Ni-based OC in the second bed. The authors studied the effects of a variation in operating temperature, pressure, and fuel composition on the process performance, concluding that power generation of such process is mainly affected by fuel type. Noorman et al.8 carried out an experimental study on the effect of operating temperature, feed oxygen concentration, and steam-to-fuel ratio on the reactivity of a Cubased OC operated in a fixed bed CLC process. The authors found that operating temperature is the 7 Environment ACS Paragon Plus


critical parameter of this process, since high temperature could endanger the long-term stability of the OC due to its relatively low melting point. Cheng et al.9 tested a number of oxygen carriers in a large laboratory fixed bed reactor using methane as fuel. They found that Cu-based OC shows both the highest reactivity and resistance to carbon deposition among all the examined carriers. Noorman et al. demonstrated both experimentally10 and numerically11 the feasibility of a CLC process in a fixed bed filled with a Cu-based OC. Moreover, Noorman et al. developed a mathematical tool to investigate CLC process in fixed bed with Cu-based OC considering both a single OC particle12 and the whole reactor13. Fernandez et al.14 developed a dynamic mathematical model in order to investigate the oxidation of Cu-based OC in an adiabatic fixed bed reactor, concluding that the use of a lean oxygen-nitrogen mixture (O2 content about 4-6%) and of a small Cu amount in the carrier (20-33%) is fundamental both to keep reactor temperature below 1200 K and to achieve highefficiency of CLC process. Although CLC in fixed bed using Cu-based OC received great attention over the last years, only few papers deal with its integration with a power plant. Hamers et al.15 evaluated the performance of a fixed bed CLC system using Cu-based OC integrated in a combined cycle power plant. Such integration poses a number of challenges, mainly associated to the need to operate several fixed bed reactors so to produce a hot gas stream (at approximately constant temperature and mass flow rate) and to ensure the highest power production of the plant by choosing the best operating conditions. This paper aims to further study such integration, by proposing a different system layout and an alternative operating strategy with respect to those reported in the above cited paper, in the view of supporting the industrial spread of such technologies. Due to the urgency of facing climate change, an increase in commercial interest towards CCS systems can be forecasted in the next few years. For this reason, a detailed economic analysis of the proposed CLC power plant is presented in terms of levelised cost of energy and cost of CO2 avoided. Eventually, the economic performances of several power plants, integrated with different CCS systems, are compared to that of the proposed layout. 2. Process layout While the oxidation of oxygen carrier is always exothermic, the reduction can be exothermic or endothermic depending on the solid material and reducing gas used. In the case of Cu-based oxygen carrier, both the carrier oxidation and reduction reactions are exothermic, obviously with different reaction enthalpies. Thus, from an energetic point of view, Cu-based CLC processes can be regarded as a four stages sequence, namely: (I) oxidation of carrier and heat generation at high temperature (Oxidation Stage, OS), (II) first heat removal (Heat Removal after Oxidation Stage, 8 Environment ACS Paragon Plus

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HROS), (III) reduction of carrier and heat generation at low temperature (Reduction Stage, RS), and (IV) second heat removal (Heat Removal after Reduction Stage, HRRS). Over the first stage, air is fed to the reactor and the OC is exothermically oxidized. At the end of OS, most of the reactor is at the maximum temperature reached during the oxidation stage (Tmax,OS)11 thus, over the second stage, air is still fed to the reactor in order to remove the heat stored during the previous stage, thus obtaining a high temperature gas stream for subsequent power generation. During the third stage, exothermic fuel oxidation occurs and, after a CO2/H2O separation step, a stream of concentrated CO2 ready for storage leaves the system. At the end of RS almost the whole bed is again at high temperature (Tmax,RS), still suitable for power generation thus, over the last stage, nitrogen is supplied to the reactor. In order to avoid the formation of a potentially explosive of air and fuel at high temperature, a Purge Stage (PS) is mandatory at the end of HROS, so that O2 concentration drops to zero along reactor before RS starts. As a consequence, the CLC process under investigation includes OS-HROS-PS-RS-HRRS that cyclically follow each other in a fixed bed reactor. Figure 1 reports a general power plant layout based on such process. Two main sections can be individuated, namely a CLC and a power one. In the first one, several adiabatic reactors should be operated in parallel in order to continuously generate two gas streams at almost constant mass flow rates and temperatures16. In the second one, the gas stream after HROS is fed to a semi-closed gas turbine using air (T1 in Figure 1) as main working fluid so that the power Pel is generated, while the gas stream after HRRS is used to warm up air at the temperature required by the process via heat exchanger HE. N2 at the outlet of OS and PS is purged out so that a make-up of fresh N2 is required. Additional air and CH4 compressors (C1 and C2 in Figure 1, respectively) are used to reach the required initial conditions for the oxidation, heat removal and reduction stages, respectively. The mixture of H2O/CO2 at the outlet of the CLC unit is sent to a condenser so to obtain a stream of concentrated CO2.



Figure 1 – General power plant layout. 3. Mathematical model 3.1 Kinetic scheme Since copper oxide on γ-Al2O3 was selected as oxygen carrier, the associated kinetic reactions scheme is reported in Table 1, along with standard enthalpies of reactions. During the oxidation stage, the oxygen carrier is oxidized by air according to reaction (R1). During the reduction stage, methane is oxidized to CO2 and H2O via reaction (R2)17. In consideration of the operating conditions taken into account, temperature during RS is well below 1175 K, thus the formation of cuprous copper Cu(I) and of its compounds was neglected11.


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Table 1 – Reactions scheme and associated standard enthalpies of reactions. Reaction

∆H0, kJ·mol-1

4CuO + CH → 4Cu + CO + 2H O

-312

R1

-178

R2

2Cu + O → 2CuO

The reaction rate of Cu-based oxygen carrier with fuel gas and air is expressed by means of the shrinking-core model for plate-like geometry under chemical reaction control17. Corresponding kinetic data are taken from Garcia-Labiano et al.18

3.2 Fixed bed mathematical model A one-dimensional pseudo-homogenous model was used to describe the dynamic behaviour of each adiabatic fixed bed reactor operated under CLC conditions. To validate the pseudohomogenous and 1D assumptions, it was verified the absence of radial concentration and temperature gradients, as well as the lack of both interphase and intra-particle concentration and temperature gradients. According to Gunn19, the influence of radial concentration and temperature gradients was evaluated through the ratio between reactor and particle diameter. Since in this work dr/dp>10020, no radial concentration and temperature gradients were considered. According to Fernandez et al.21, the effects of interphase concentration and temperature gradients were considered by evaluating the gas-solid mass (hm) and temperature (km) transfer coefficients, respectively, by means of Eqs. (1-2)22:

ℎ = 0.357 . / ! " = 1.37

(1)

! %&' (' )' * . /

.$

(2)

Taking into account the parameters used in this work, minimum values of hm and km of about 0.05 m∙s-1 and 0.1 W∙m-2∙K-1 are achieved, respectively. Thus, due to rapid gas-solid mass and heat transfers occurring, both the interphase gradients can be safely neglected23,24. As reported by25, the satisfaction of Weisz-Prater criterion26 should be verified for each reactant and corresponding reaction in order to assess the influence of intra-particle concentration gradient:

+,- =

6 ./01 ∙31 ∙ 1 ∙45 78,: ∙;:6

(3)



in which the observed reaction rate (robs) deviates less than 5% from the rate calculated assuming a homogenous mixture27 and the effective diffusion coefficient of specie i (De,i) was calculated according to25 and reference therein. It was observed that NWP is always ? =>? = =>? + )' = `6 QA, 0R = >b6 QA, 0R = >;` QA, 0R = 0 Z > QA, 0R = > [ Z Y

;

;, 4C

GQA, 0R = G NQA, 0R = N?S

(5)

The initial conditions of subsequent stages were assumed to be equal to the last computed values from the previous one, according to the cyclic nature of CLC process. Further details on the mathematical expression of cyclic initial conditions can be found in6. For the evaluation of temperature, pressure and composition dependencies of reaction enthalpies, transport coefficient and gas properties, state-of-the-art correlations and assumptions were adopted from Han et al.25 and reference therein. The mathematical model was solved using the commercial software package Comsol Multiphysics®. Reactor length (L) was discretized with 500 nodes and it was verified that further refinements of the spatial mesh do not produce any appreciable changes in the calculated temperature and concentration profiles. The validation of above-described mathematical model against literature data11 was discussed elsewhere31. 4. Results and discussion Integration results for the cyclic process after a number of OS-HROS-PS-RS-HRRS stages are discussed in Section 4.1. Starting from these outcomes, a potential strategy for the operation of several reactors in parallel to obtain a continuous production of high temperature gas streams after both HROS and HRRS is developed in Section 4.2. Eventually, the integration of proposed CLC system and stationary power plant is evaluated by an energetic and economic point of view in Sections 4.3 and 4.4, respectively.



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4.1 CLC process in single fixed bed The model parameters used to simulate the CLC process (Tab. 3) were taken from Noorman et al.11, with the exception of gas feeding pressure (Pin), initial bed temperature (T0) and inlet gas temperature (Tin). Pin was set to 2.0 MPa, a typical value for natural gas combined cycle power plants36. T0 value should be set higher than activation temperature of Cu reduction reaction but, at the same time, as low as possible in order to avoid irreversible catalyst thermal deactivation. T0 equal to 850 K was chosen as a good compromise. Indeed, the selected value of T0 is adequately close to the reduction reaction threshold temperature (≈873 K18). Also, according to Eq. (6)11, maximum reachable reactor temperatures during OS and RS are 1175 K and 1045 K, respectively. According to4, Cu-based OC is still able to work at such temperatures (melting point ≈1400 K18).

∆GI = e51∙fgeh

∆]d

ij geh ∙0d

k

e5l,∙fl,

(6)

il,

Since the gas feed temperature during heat removal stages corresponds to the initial bed temperature for subsequent reduction and oxidation stages, Tin was set at 850 K too. Assuming that the process starts when the OC is in its fully reduced forms (elemental Cu), a single CLC cycle consists of OS-HROS-PS-RS-HRRS stages, as described in section 2. The details of the boundary conditions associated with each stage are reported in Table. 4. The gas mass fluxes (Gs) during OS and RS were taken from11; in order to increase power production, a higher Gs was adopted during both HRRS and HROS, which allowed reaching a reasonable CLC period21 and a gas superficial velocity of about 1.5 mˑs-1, that is close to the normal range of operation for chemical looping processes in fixed bed32. The period of each stage was not fixed a priori, but a controller automatically sets it. To clearly explain the adopted switching strategy, Fig. 2 reports monitored outlet gas concentration (a) and temperature (b) during the first CLC cycle. In detail: i) the controller dictates the switch between OS and HROS when outlet O2 concentration reaches approximately its inlet value (blue zone in Fig. 2), in order to fully oxidise the OC in the reactor, ii) the subsequent HROS is continued until outlet gas temperature drops below 1175 K (first white zone in Fig. 2), in order to feed a gas stream at approximately constant temperature to a downstream unit, iii) the subsequent PS is extended until outlet gas temperature achieves its inlet value (red zone in Fig. 2), in order to restore a flat T0 temperature profile along the whole reactor, iv) the switch between RS and HRRS is dictated when outlet CH4 concentration reached its inlet value (yellow


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zone in Figure 2), in order to fully reduce the carrier in the reactor, v) the last HRRS stage was carried out until a flat T0 profile was restore in the whole reactor (second white zone in Figure 2). Table 3 – Model parameters. Parameter

Value

Pin, MPa

2.0

Tin, K

850

T0, K

850

L, m

1.0

dr, m -1

cps, J∙kg ∙K

0.5 -1

1080

ρs, kg∙m-3

3000

dp, m

2.0∙10-3

wact0

0.2

CCu,max, mol∙m-3

4460

εs

0.4

Table 4 – Detailed boundary conditions. OS

HROS

PS

RS

HRRS

Tin, K

Tin,CLC

Tin,CLC

Tin,CLC

Tin,CLC

Tin,CLC

Pin, MPa

2.0

2.0

2.0

2.0

2.0

CN2,in=(Pin/R/Tin)

CCH4,in=(Pin/R/Tin)

CN2,in=(Pin/R/Tin)

Ci,in, mol∙m-3 GS,in, kg∙m-2∙s-1 3 -1

Qin, m s

CO2,in=0.21(Pin/R/Tin)

CO2,in=0.21(Pin/R/Tin)

CN2,in=0.79(Pin/R/Tin)

CN2,in=0.79(Pin/R/Tin)

1.0

4.0

1.0

0.1

4.0

0.02

0.08

0.02

0.004

0.08



Figure 2 – Monitored outlet gas concentration (a) and temperature (b) during the first CLC cycle (OS in blue, PS in red, RS in yellow, HROS and HRRS in white).

Several CLC cycles were simulated, verifying that regime conditions were attained after just the first cycle. Indeed, for the adopted control policy, the temperature profile along the bed at the end of each CLC cycle is equal to the initial one, i.e. T0 along the whole reactor. As a consequence, the figure corresponding to several CLC cycles were not reported for the sake of brevity. From the above, it can be inferred that the CLC period is 3500 s (ΠCLC). Table 5 reports the period of each stage.

Table 5 – Period of each stage. Stage

Period, s

OS

312

HROS

1600

PS

308

RS

230

HRRS

1050

Total

3500

4.2 CLC process in a fixed bed reactor network By analysing the periods reported in Tab. 5, it is clear that several reactors in parallel have to be operated under above described CLC conditions in order to continuously produce two hot gas streams at approximately constant temperature during both HROS and HRRS. Therefore, the 16 Environment ACS Paragon Plus

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required minimum number of the reactor was assessed by the ratio between the period of whole CLC cycle and the minimum one between the HROS and HRRS, as reported in Eq. (7):

+ ?S = mn o ?SQp

pqrq

stuv ,psttv R

w

(7)

where the ceil function rounds Nmin up to the nearest integer. From Eq. (7), Nmin equal to 4 was obtained. This number would be increased only if the pressure drop over each stage goes above inlet value by 15%33. According to the approach described by Diglio et al.34, it was checked that the pressure drop in each stage is at worst 8% of inlet value. Thus, 4 reactors (Nr = Nmin) in parallel were chosen to carry out the target process. It is noteworthy that, as previously described, HRRS was carried out until a T0 temperature flat profile is attained along the whole bed in order to restore initial conditions for the subsequent OS. However, T quickly decreases towards T0 over the last HRRS time steps (see Figure 2), thus making the gas stream too cold to keep pre-heating air gas stream at the suitable temperature required by the process. As a consequence, the gas stream leaving each reactor at an outlet temperature below 1020 K (i.e. from t=914 s up to the HRRS end) cannot be sent to HE. The operation strategy of reactor network is to trigger on the reactors one after another, with a proper displacement time (tdisp). In order to reduce the mass flow rate fluctuation at the inlet of the gas turbine, tdisp was selected with the aim of limiting the variation over time of the number of reactors working in HROS, respectively. Figure 3A reports the percentage of the whole CLC period over which j reactors (j=1,2,3) work simultaneously in such stage (ν) along tdisp: ν reaches a maximum in the time-frame 777÷914 s, when 2 reactors are in HROS (red line) for about 77% of the whole CLC period. It is important to note here that ν, and consequently the operation strategy, is not affected by the initial stage chosen to start the process. As a consequence, in this work tdisp was arbitrarily set equal to 904 s. It is important to underline here that, for any given displacement time, at least one of the network reactors should be in HRRS in order to be regenerated in the subsequent cycle stage. Figure 3B reports the number of the reactors working in HROS over time when tdisp=904 s. A constant mass flow rate, equal to (Nr,HROS-α+β)∙GS,HROS, is fed to turbine system during HROS. Therefore, a thermal buffer is used to level the mass flow rate fluctuation. In particular, the mass flow rate at the outlet of HROS is split by sending α part to thermal buffer and the remaining (1-α) part to turbine system. The thermal buffer supplies the turbine system with a mass flow rate equal to

β∙GS,HROS. The values of α and β over time are reported in Tab. 6. 17 Environment ACS Paragon Plus


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Table 6 – α and β values over CLC period. Time

Reactors in HROS (Nr,HROS)

α

β

(Nr,HROS-α+β )

0.94∙ПCLC

2

0.23

0

1.77

0.06∙ПCLC

1

0

0.77

1.77

Figure 3 – Percentage of the whole CLC period (ν) in which 1 (blue line), 2 (red line) or 3 (green line) reactors work simultaneously in HROS by varying tdisp (A) and number of reactors working in HROS over time when tdisp=904 s (B).


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Figure 4 – Outlet temperatures of the streams made up by mixing the gaseous leaving the reactors working in HRRS (solid black line) and HROS (dash black line); H2O (dot red line) and CO2 (dashdot red line) molar fractions in the stream leaving the only reactor working in RS.

Figure 4 reports the main outcomes of the reactor network, i.e. outlet temperatures of the streams made up by mixing the gas leaving the reactors working in HRRS (solid black line) and HROS (dash black line), H2O (dot red line) and CO2 (dash-dot red line) molar fractions in the stream leaving the reactors working in RS. Considering that the reactors in CLC section start from HRRS, the streams from HROS and RS were detected at the exit of reactor network only after 1362 s (ПHRRS+ПOS) and 3270 s (ПHRRS+ПOS+ПHROS+ПPS), respectively, i.e. the time required by the first trigged on reactor to reach the first heat removal after oxidation and reduction stages, respectively. Of note, by operating 4 reactors in parallel under CLC conditions, it is possible to obtain two hot gas streams at approximately constant temperature of about 1175 K and 1045 K, respectively, for subsequent turbine expansion, with the important feature that a stream of concentrated CO2 can be obtained after condensation process. 4.3 Integration between the proposed CLC process and a stationary power plant: Energetic Analysis The overall process layout is depicted in Figure 5. CLC section consists of a network of 4 fixed bed reactors, while power section comprises a gas turbine system. Regarding CLC section, the 19 Environment ACS Paragon Plus


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inlet gas of each reactor is air, nitrogen or methane according to the current stage. Nitrogen from collector HRRS warms inlet air for OS and HROS via heat exchanger HE. N2 coming out from HE is at suitable temperature to be directly re-used as feed stream for HRRS and PS. Gas stream from collector HROS is fed to a thermal buffer, while H2O and CO2 from collector RS warms methane via HE1. Outlet N2 from OS and PS, along with that coming from the last time steps of each HRRS cycle, are purged out (N2 purge in Fig. 5). Thus, a continuous make-up of fresh N2 is required. Since the mass flow rate of nitrogen make-up is 0.05% of that re-circulated coming from HRRS, i.e. around 1000 m3·yr-1, a fresh make-up at ambient temperature does not alter sensibly the temperature of recirculating N2.

Figure 5 – Layout of the SE-SMR process integrated with stationary power plant.

The net efficiency of the whole process can consequently be evaluated as: --,;x; =

-8y,z8h

{1,tv |l x]}qs^

=

-8y,stuv %∑6z -8y,e/,z * {1,tv |l x]}qs^


(8)

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The estimation of the power generated by the turbine T1 (Pel,HROS), as well as the power demand of compressors (PC,i) is described in Supporting Information S1 in detail. The results of the energetic analysis are reported in Tab. 7.

Table 7 – Results of the energetic analysis. Pel,HROS, MW

1.21

Pel,com,1, MW

0.71

Pel,com,2, MW

1.64∙10-2

Pel,net, MW

0.50

enet

51%

The net power generation of the system is about 0.50 MW, with a net efficiency of 51%, while its net specific power generation is about 7.08·10-3 MW∙kg-1. Comparison with the efficiency and specific power generation of several benchmark technologies is presented in Tab. 8. In particular, both power plants with and without a carbon capture and storage system were considered. The latter are an Integrated Gasification Combined Cycle (IGCC) power plant, a SuperCritical-Pulverized Coal (SC-PC) power plant and a Natural Gas-fired power plant (NG). The former are an IGCC based on a pre-combustion CCS in which a water gas-shift reactor enrich H2 content of syngas deriving from coal gasification and a CO2 adsorption process separates CO2 from H2, this latter is used for turbine expansion (IGCC with pre-combustion CCS); a SC-PC in which pulverized coal firing technology (steam temperature approximately 870 K, pressure around 1617 MPa35) is integrated with a chemical absorption process in aqueous amines to capture CO2 from the exhaust of combustion process (SC-PC with post-combustion CCS). It is clear from Tab. 8 that among the CCS based power plants, the proposed system shows the highest performance, which is also very close to that of power generation system using similar CCS technology found in literature36.



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Table 8 – Net efficiency of several power plants. Plant

Net Efficiency, %

Specific Power,

Reference

MW∙kg-1 Without CCS IGCC

45.2

3.77·10-3

37

SC-PC

44.1

3.68·10-3

37

NG

55.0

7.64·10-3

38

With CCS IGCC with pre-combustion CCS

35.3

2.94·10-3

37

SC-PC with post-combustion CCS

35.2

2.93·10-3

37

Proposed CLC based Power Plant

51

7.08·10-3

-

4.4 Integration between the proposed CLC process and a stationary power plant: Economic Analysis The economic performance evaluated in this work are the Levelised Cost Of Energy (LCOE) and the cost of CO2 Avoided (AC)39: T>S =

> =

;_z ∙;k`z ;z ∙-8y,z8h

x;`,qrq x;`t Xqu6,t Xqu6 ,,qrq

+ S +

;z

Xz8h,z

= NN, >T> M

(9) (10)

An existing Natural Gas-fired Power Plant (NG-PP) without CO2 capture was assumed as reference system. Since the proposed system is a near-zero CO2 emission process, eCO2,PP,CLC was assumed equal to 0. The Fixed Charge Factor (FCF) was evaluated as39


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> =

-_

aQ-_kaR

(11)

It was considered that the Project Interest Rate (PIR) and amortisation years (γ) are 8.78% and 25 years, respectively40, for both CLC power plant and reference system. The Total Capital Requirement of the proposed system (TCRPP,CLC) was assessed considering the cost of reactor network, i.e. high-temperature valves, reactors and solid material, the cost of gas turbine and compressors, heat exchangers, condenser, and thermal buffer. The detailed cost analysis of the proposed system is reported in Supporting Information S2. The Fixed Operating Maintenance costs (FOM) were assumed to be 1% of TCR for both reference and proposed system39. The Variable Operating Maintenance costs (VOM) were assumed to be 4% of TCR for the reference system, while, as reported by37, in the case of proposed system VOM were evaluated as: = >G + + >G = = =

(12)

;∙%.∙a *∙%{1,tv ∙|l ∙qu6,/h *∙; a ∙;

-8y,z8h

$∙;∙ ∙-8y,z8h

(13) (14)

a ∙;1/y

$∙;∙1 ∙-8y,z8h

(15)

All parameters used for the economic analysis were reported in Tab. 9.



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Table 9 – Parameters used for the economic analysis. Parameter

Value

Reference

Net power generation (Pel,net)

0.50 MW

-

Net efficiency of proposed system (enet,PP,CLC)

0.51

-

Net efficiency of reference system (enet,REF)

0.55

38

Specific cost of reference system (CREF)

700∙103 €·MW-1

38

Total Capital Requirement for reference system CREF∙Pel,net

-

(TCRREF) Capacity Factor for reference system (CFREF)

0.50

38

Capacity Factor for proposed system (CFPP,CLC)

0.70

37

Specific cost of natural gas (SFCREF)

1.3 €·MWh-1

38

Specific cost of CH4 (SFCPP,CLC)

1.9 €·MWh-1

38

CO2 transport and storage cost (CTS)

10 €·tonCO2-1

37

Oxygen carrier life time (lts)

5y

37

High-temperature valve life time (ltv)

10 y

40

CO2 emission reference system (eCO2,REF)

0.379 tCO2·MWh-1

38

Figure 6 reports the contribution of the main components on the total capital requirement of proposed system. The reactor network is the most expensive component, which accounts for the 58% of the whole TCR. In particular, the cost of high-temperature valves is preeminent, in accordance with the data reported by41. This means that the operating conditions of CLC process should be carefully checked in order to increase as much as possible the whole CLC period in order to decrease the number of reactors working in parallel, consequently also the number of valves required too. It is noteworthy that the contribution of thermal buffer is negligible to the total capital requirement.


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Figure 6 – Relative cost of the proposed CLC Power Plant.

The resulting economic performances are reported in Table 10. The LCOE of the reference system is very close to literature data for natural gas-firing power plant38. The proposed system shows better economic performance than a conventional natural gas-firing power plant with postcombustion CO2 capture system, where typically an organic solvent such as monoethanolamine is employed38, for which LCOE is around 80 €·MWh-1 and AC is about 80 €·tonCO2-1 according to42. Moreover, the proposed system has also better economic performance than CCS technologies reported in Tab. 8. Indeed, in the case of IGCC with pre-combustion CCS, LCOE is 110 €·MWh-1 and AC ranges from 30 €·tonCO2-1 to 80 €·tonCO2-1 according to the reference system considered, while in the case of SC-PC with post-combustion CCS, LCOE is 102 €·MWh-1 and AC varies from 55 €·tonCO2-1 to 70 €·tonCO2-1 depending on the reference system considered37. Instead, as reported by Erans et al.43, a natural gas-firing power plant integrated with calcium looping is characterised by LCOE (44 €·MWh-1) better than that of proposed system and a similar AC (30 €·tonCO2-1).



Page 26 of 34

Table 10 – Economic performance. Index

Value

Levelised Cost Of Energy reference system (LCOEREF)

42 €·MWh-1

Levelised Cost Of Energy proposed system (LCOEPP,CLC)

54 €·MWh-1

Cost of CO2 Avoided (AC)

31 €·tonCO2-1

A sensitivity analysis was carried out by varying the cost of thermal buffer, carbon transport and storage, fixed beds, heat exchangers, condenser, gas compressor and turbines, high-temperature valve, oxygen carrier, and fuel by ±15%. Figure 7 reports the sensitivity analysis in terms of LCOE, while Figure 8 shows AC. The price of fuel has the strongest impact on both LCOE (±11%) and AC (±50%). This is in agreement with literature data40. It should be taken into account that this is the most market-depending cost and its value in the next years is the most uncertain. It was verified that the cost of N2 can be neglected since the small amount of fresh make-up assessed, accounting for a mere 500 €·yr-1. According to the results showed in Figure 6, the high-temperature valve is the most expensive component of the whole proposed system, thus its impact on LCOE (±3%) and AC (±12%) is much higher than the one of the other components. Therefore, a further sensitivity analysis was assessed by varying high-temperature valve lifetime. The results in terms of LCOE (black line) and AC (red line) are reported in Figure 9. By inspecting this figure it is clear that the valve lifetime is the key parameter for the economic feasibility of the proposed system.


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Figure 7 – LCOE sensitivity analysis.

Figure 8 – AC sensitivity analysis.



Page 28 of 34

Figure 9 – LCOE and AC varying high-temperature valve lifetime.

The economic analysis of the proposed system was completed by assessing a discounted cash flow analysis. In particular, the Cumulative Discounted Cash Flow (CDCF) was evaluated as37: qrq >B> = ∑ S Qak7_Rz

;|

(16)

where a discount rate (DR) of 8% is used and discount starts from the first day of plant operation. The construction period was set to 3 years: the cash flow of CLC (CAFCLC) power plant was imposed equal to the 30% of TCRPP,CLC in the 1st construction year, 45% in the 2nd year and 25% in the 3rd year. Instead, during the remaining 25 operation years the cash flow of the proposed system was evaluated as the difference between the sell price of energy and levelised cost of energy. To account for initial inefficiencies, this last one was assessed by considering a capacity factor of 65% in the first operation year and 85% in the second operation year, while for the remaining 23 years of operation the maximum capacity factor was taken into account37. Figure 10 shows the CDCF along the 25 operating years considering a base case (110 €·MWh-1 according to37) and by varying the sale price of energy of ±15% with respect to the base case.


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Figure 10 – Cumulative discounted cash flow for an electric price of 110 €ˑMWh-1 (base case, solid line) and considering a variation of -15% (dash line) and +15% (dash-dot line) of electric price with respect to the base case.

In the base-case, the pay-back period is about 2.8 years, with a cumulative profit over 25 years about 5 times greater than the total capital requirement (1.02 M€ and 205 k€, respectively). Instead, if the sell price of energy increases of 15% with respect to the base-case, the investment become more profitable (≈1.6 M€) and a faster pay-back is achieved (≈2.4 years). On the contrary, if this price is lower than 15% when compared to the base-case, the investment appears to be very risky. It is noteworthy that the same analysis was carried out by varying of ±15% the CTS with respect to the value reported in Tab. 8 and its impact on the base-case is negligible, thus the corresponding figure was not reported for the sake of brevity. 5. Conclusions A small scale power generation system with inherent CO2 capture was proposed and numerically investigated. The developed layout consists of two sections: a CLC one where, using Cu as oxygen carrier and methane as fuel, four adiabatic fixed beds work in parallel in order to continuously produce two hot gas streams, one for turbine expansion and another one to pre-heat



feeding gas stream, and a power section devoted to power generation through combined cyclebased gas turbine. Operating conditions of CLC process in fixed bed were investigated by means of a onedimensional mathematical model. CLC process can be seen as a sequence of five stages, namely oxidation, heat removal after oxidation, purge, reduction, and heat removal after reduction, that periodically follow each other. The period of each stage was automatically controlled by monitoring outlet gas concentrations and temperature. The adopted controlling strategy aims to completely oxidize and reduce the oxygen carrier during oxidation and reduction stages, respectively. Under the above conditions, two hot gas streams can be produced at approximately constant temperature of about 1175 K and 1045 K, respectively. The first one is used for turbine expansion while the heat content of the second one to warm up the fresh air feed. A proper displacement time between the triggering of each reactor and a thermal buffer are required to level the mass flow rate before the expansion into turbine system composing the power section. The calculated net power generation was about 0.50 MW with a global energy efficiency, higher than that of alternative CCS technologies, of 51%. The economic analysis of the proposed power generation system revealed that the reactor network is the most expensive component (58%), almost entirely due to the high-temperature valve (81%). The levelised cost of electricity and CO2 avoided are 54 €·MWh-1 and 31 €·tonCO2-1, sensibly lower than those of alternative CCS technologies. The sensitivity analysis showed that the economic performance of the proposed system is mainly affected by the specific cost of fuel. Also, the lifetime of high-temperature valve is the key parameter for the economic feasibility of this system. A discounted cash flow analysis demonstrated that in the base-case (sell price of energy 110 €·MWh-1) the pay-back period is about 2.8 years, with a cumulative profit over 25 years about 5 times greater than the total capital requirement. When the energy sell price is increased by 15% with respect to the base-case, a faster pay-back time is achieved (≈2.4 years) and the investment become more profitable (≈1.6 M€ over 25 years). Instead, the investment appears to be very hazardous considering a sale price of energy lower than 15% with respect to the base-case. Supporting Information Energy (S1) and cost analyses (S2) are discussed in detail. This material is available free of charge via the Internet at http://pubs.acs.org.


Page 30 of 34

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Techno-Economic Evaluation of a Small-Scale Power Generation Unit

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