ECRA Chair From CO to Energy: Carbon Capture in Cement Production and ... · to Energy: Carbon...

ECRA Chair “From CO2 to Energy:

Carbon Capture in Cement Production and its Re-use”

Third Annual Report

Period: May 2015 - April 2016

http://hosting.umons.ac.be/html/ecrachair

Contributors to the report:

Prof. Diane Thomas, Chemical & Biochemical Process Engineering Unit, UMONS Academic Coordinator of the Chair Prof. Guy De Weireld, Thermodynamics Unit, UMONS Academic Supervisor of the Chair Dr Lionel Dubois, Chemical & Biochemical Process Engineering Unit, UMONS Scientific Coordinator of the Chair Ir Sinda Laribi, Chemical & Biochemical Process Engineering Unit, UMONS PhD Student Ir Seloua Mouhoubi, Chemical & Biochemical Process Engineering Unit, UMONS PhD Student Ir Nicolas Meunier, Thermodynamics Unit, UMONS PhD Student Ir Remi Chauvy, Thermodynamics Unit, UMONS PhD Student

Acknowledgments:

ECRA is warmly acknowledged for the financial support of the Academic ECRA

Chair at the Faculty of Engineering at UMONS (Belgium).

We would like also to acknowledge the HeidelbergCement Group for the funding accorded to the ECRA Chair.

ECRA CHAIR – Third Annual Report (May 2015 – April 2016)

Table of contents

1. STATUS OF THE ECRA CHAIR ACTIVITIES .......................................................................... 6

1.1. FRAMEWORK OF THE ECRA CHAIR ............................................................................................................... 6 1.1.1. PhD Theses ...................................................................................................................................... 6 1.1.2. Association of students works ....................................................................................................... 10 1.1.3. Post-doc works .............................................................................................................................. 10

1.2. PROLONGATION OF THE ECRA CHAIR AT UMONS ....................................................................................... 11 1.3. SCIENTIFIC COMMITTEE MEETINGS ECRA-UMONS ...................................................................................... 12

1.3.1. Meeting (ECRA – UMONS) at Mons on 22th October 2015 ............................................................ 12 1.3.2. Meeting (ECRA – UMONS) at Lixhe on 13th April 2016 .................................................................. 13

1.4. ECRA PRIZE........................................................................................................................................... 13 1.5. ECRA CHAIR WEBSITE.............................................................................................................................. 14

2. ADVANCES IN THE PHD THESIS OF NICOLAS MEUNIER .................................................. 15

2.1. ADVANCES IN THE DEHYDRATION UNIT ....................................................................................................... 15 2.2. ADVANCES IN THE METHANOL CONVERSION PROCESS .................................................................................... 15

2.2.1. Description of the process flowsheet ............................................................................................. 15 2.2.2. Results and discussion ................................................................................................................... 17 2.2.3. Estimation of the operating and capital costs of the installation ................................................. 19 2.2.4. Description of the micro-pilot installation ..................................................................................... 22 2.2.5. Conclusion and perspectives.......................................................................................................... 24

3. ADVANCES IN THE PHD THESIS OF SINDA LARIBI ........................................................... 25

3.1. PARTIAL OXYFUEL COMBUSTION CAPTURE .................................................................................................... 26 3.1.1. Absorption-regeneration process .................................................................................................. 26 3.1.2. Experimental tests conducted at lab scale .................................................................................... 27 3.1.3. Micro-pilot tests for the best solvents screened ............................................................................ 29 3.1.4. Estimation of the total energy required under partial oxyfuel conditions..................................... 32

3.2. FULL OXYFUEL COMBUSTION CAPTURE ......................................................................................................... 33 3.2.1. SOx/NOx reactions chemical mechanism ...................................................................................... 33 3.2.2. Simulation results of the Sour-Compression Unit (SCU) absorption performances ....................... 34

4. ADVANCES IN THE PHD THESIS OF REMI CHAUVY ......................................................... 36

4.1. METHODOLOGY AND DEVELOPMENT FOR THE PRE-SELECTION .......................................................................... 37 4.1.1. Definition of the criteria ................................................................................................................ 37 4.1.2. Identification of the CO2 conversion routes and their level of maturity ........................................ 38 4.1.3. Size of the CO2 utilization .............................................................................................................. 40

4.2. EVALUATION FOR THE PRE-SELECTION ......................................................................................................... 42 4.3. METHODOLOGICAL SELECTION ................................................................................................................... 43

4.3.1. Definition of the criteria and indicators ........................................................................................ 43 4.3.2. First output and outlook ................................................................................................................ 45

5. BEGINNING OF THE PHD THESIS OF SELOUA MOUHOUBI ............................................. 46

5.1. THESIS CONTEXT AND CONTENT.................................................................................................................. 46 5.2. LIQUID-LIQUID PHASE CHANGE SYSTEMS ...................................................................................................... 46

5.2.1. Thermomorphic lipophilic amine solvents ..................................................................................... 47 5.2.2. DMX Solvents................................................................................................................................. 48 5.2.3. Self-concentrating solvent ............................................................................................................. 48

ECRA CHAIR – Third Annual Report (May 2015 – April 2016)

5.2.4. {MAPA+DEEA} blend ...................................................................................................................... 49 5.2.5. {BDA+DEEA} blend ......................................................................................................................... 49 5.2.6. {TETA+DEEA} blend ........................................................................................................................ 50

5.3. SUMMARY AND WORK PROSPECTS .............................................................................................................. 50

6. ADVANCES IN THE POST-DOCTORAL WORKS OF LIONEL DUBOIS.................................. 51

6.1. COMPLEMENTARY STUDIES ON CO2 MINERALIZATION INTO OLIVINE-DERIVED COMPOUNDS ................................... 51 6.1.1. pH-swing mineral carbonation process ......................................................................................... 51 6.1.2. Cement wastes utilization in CO2 mineral carbonation processes................................................. 54

6.2. COMPARISON OF VARIOUS CONFIGURATIONS OF THE ABSORPTION-REGENERATION PROCESS FOR THE POST-COMBUSTION

CO2 CAPTURE APPLIED TO BREVIK CEMENT PLANT FLUE GASES ........................................................................................ 58 6.2.1. Introduction ................................................................................................................................... 58 6.2.2. Simulation of different process configurations ............................................................................. 58 6.2.3. Results and discussion ................................................................................................................... 63 6.2.4. Conclusions and perspectives ........................................................................................................ 67

7. GENERAL CONCLUSIONS AND PERSPECTIVES ................................................................ 69

7.1. GENERAL CONCLUSIONS OF PROGRESS OF THE ECRA CHAIR SCIENTIFIC ACTIVITIES ............................................... 69 7.2. ECRA CHAIR SCIENTIFIC ACTIVITIES FOR THE NEXT YEAR .................................................................................. 70 7.3. EXTERNAL COMMUNICATION ..................................................................................................................... 71

7.3.1. Publications ................................................................................................................................... 71 7.3.2. Future planned communications ................................................................................................... 74

7.4. GLOBAL PERSPECTIVES OF THE ECRA CHAIR ................................................................................................. 75

8. REFERENCES .................................................................................................................... 76

9. ANNEXES ......................................................................................................................... 78

ECRA CHAIR – Third Annual Report (May 2015 – April 2016) 6

1. STATUS OF THE ECRA CHAIR ACTIVITIES

The main objective of the ECRA Chair at UMONS, launched in April 2013, is to create a centre of scientific expertise in the specific field of “Carbon capture in cement production and re-use” and to promote research and innovation in the topic through the achievement of different activities as listed in Figure 1.

Figure 1: “Spirit” of the ECRA Chair

Most of them have been developed as evidenced by this report.

1.1. Framework of the ECRA Chair

The different ECRA Chair activities at UMONS are at this date supported by: - four PhD students: N. Meunier, S. Laribi, R. Chauvy and S. Mouhoubi; - one postdoctoral researcher: L. Dubois; - professors of the Faculty of Engineering of UMONS: D. Thomas, G. De Weireld and P.

Lybaert.

1.1.1. PhD Theses

1°) The scientific studies were launched at the beginning of the academic year 2013-2014 (in September 2013) through two PhD theses co-funded by ECRA and UMONS (Special Fund for Research / Research Institute for Energy). These were recently (during this academic year 2015-2016) completed by two more PhD subjects. We can present schematically the different and consistent parts of the research works as follows (see Figure 2), with two main subjects: CO2 capture and purification, and CO2 conversion. All these will be more developed later in this document.


Figure 2: Research subjects considered into the ECRA Chair

- The first thesis is carried out by Nicolas Meunier (Master’s degree in Chemistry/Material Science from UMONS and TU Wien). He began on 1st September 2013.

Nicolas Meunier works on a subject entitled “CO2 capture in cement production and re-

use: optimization of the overall process”.

N. Meunier has obtained a research grant (beginning in October 2014) of 2 years renewable once by the FNRS (Belgian National Fund for Scientific Research) for carrying out his PhD thesis on the CO2 catalytic conversion into methanol. The main objectives of this research are to perform a multi-scale characterization of the transfers of the different compounds for the catalytic reaction of CO2 conversion into methanol (i.e. CO2, hydrogen, water, and methanol) and to measure the influence of gaseous compounds (O2, SOx and NOx) on CuO/ZnO/Al2O3-type catalysts performances, stability and lifetime. To this end, laboratory and pilot scale units will be established and experiments will be performed and then modelled by homemade mathematical codes to determine the different kinetic laws and mechanisms of the reactions considered for the CO2 conversion into methanol. The influence of some promoters on performances, stability and lifetime of the CuO/ZnO/Al2O3-type catalyst used for the CO2 conversion into methanol will also be investigated. Moreover, predictive simulations (on Aspen Plus®) will also be performed in order to verify the relevance of the kinetics and mechanisms models and to set the minimal CO2 purity required for the CO2 conversion into methanol according to user-defined methanol purity. Finally, the research will propose an integrated process with optimized operating parameters and considering energetic and cost integration for the design and sizing of the installation. The

CO2 conversion


subject of his thesis will address the global (capture/ purification/ conversion) process chain and the specificities of the cement industry. In his thesis N. Meunier also considers the simulation of the dehydration and cryogenic units as parts of the CO2 Purification Unit (CPU) applied to carbon dioxide coming from oxy-fuel cement plants.

- The second thesis began on 1st January 2014. Miss Sinda Laribi (Applied License in Industrial Chemistry (ESSTunis) and Master of Chemical Engineer (ENIGabes)) was recruited as a PhD student.

She works on “Capture and purification processes applied to CO2 derived from cement

industry for conversion into methanol”.

Both for post-combustion and oxy-fuel combustion CO2 capture processes, the outcoming CO2 flux must be purified in order to be re-used, especially in the methane or methanol process. The purpose of the PhD thesis is to review, understand and simulate (allowing capture and energy performances estimations) the different flue gas treatments needed for the post-combustion CO2 capture and for the oxy-fuel combustion capture (rich CO2 flow purification). S. Laribi focuses on two main subjects, namely the simulation of the Sour Compression Unit

(SCU) as part of the CO2 Purification Unit (CPU) applied to carbon dioxide coming from oxy-

fuel cement plants, and the experimental and simulation study of CO2 absorption-

regeneration process into amine-based solvents applied to conventional flue gases or, from a

more innovative point of view, flue gases coming from O2-enriched cement plants (“hybrid

CO2 capture process”).

Regarding the first point detailed investigations and bibliographic researches are required to

implement the correct reaction mechanism, involved in the SOx and NOx reactive absorption

process, with the adequate characteristics (equilibrium and kinetic constants) in Aspen Plus

software in order to obtain realistic simulation results.

Concerning the second point, absorption test runs were carried out thanks to a laboratory gas-

liquid contactor to test several amine-based solvents and other solvents such as hybrid ones

(combination between an amine and another compound such as an acetal) to test their

absorption performances at high CO2 contents (up to 60%) into the flue gas. These operating

conditions are related to the possible interest for the cement industry to consider a

combination of O2-enriched conditions (with or without gas recirculation) and a post-

combustion CO2 capture process. This screening led to the selection of solvents which are, in

a second step, tested in an absorption-regeneration micro-pilot- unit. Simulation tests of the

process are also carried out (micro-pilot and industrial scales) for conventional amines

(available in Aspen HysysTM).


The first CO2 capture processes investigations are then made on the amine absorption-

regeneration process, but a second technology (Pressure Swing Adsorption) is also considered.

- The third thesis began on 1st September 2015. Mr Remi Chauvy (Degree of Engineer of the University of Technology of Troyes (UTT) in Material Sciences (specialization in environmental assessment) and Master of Sustainable and Environmental Management and Engineering of UTT) was recruited as a PhD student.

He works on “Study of the potential of different CO2 conversion options for the

application of Carbon Capture and Re-use to the cement industry: simulation and technico-economic analysis”.

The purpose of this PhD thesis is to evaluate, in parallel of the CO2 conversion routes into renewable methanol already considered into the ECRA Chair, the potential of other options for the re-use of CO2 captured from cement plants. More precisely, the PhD thesis will include the following tasks: - detailed bibliographic review on the different CO2 conversion options (synthetic fuels,

chemical building blocks, mineralization, etc.) and on the optimization methods previously

applied to other industries (for example power plants) and to other CO2 valorization

methods;

- methodological selection of maximum 2 or 3 more interesting CO2 conversion routes based

on: technical aspects, economical aspects, energetical aspects, environmental aspects,

market considerations, and also taking into account the specificities linked to the cement

industry such as the CO2 quantity to convert;

- simulation (using adequate tools) of one or several CO2 conversion route(s) in order to

obtain relevant technico-economical-environmental indicators for comparison and

optimization purposes.

Different tasks are currently already investigated, giving first very interesting results.

- The forth thesis began on 1st February 2016. Miss Seloua Mouhoubi (Master in Chemical Process Engineering, University of Béjaia (Algeria) and Master in Chemical Engineering, University of Paris 6/Chimie Paris Tech/CNAM (France)) was recruited as a PhD student. She works on “Development of a simulation model of the post-combustion CO2 capture process by absorption-regeneration using demixing solvents: application to cement flue gases”.

The present subject will especially include the development of a simulation model allowing the estimation of the energy savings thanks to the demixing phenomenon. More precisely, her PhD thesis will include the following tasks: - detailed bibliographic review on the different phase change CO2 capture processes and more specifically on the demixing solvents technology;


- simulation (using simulation tools such as Aspen PlusTM or Aspen HysysTM) of the post-combustion CO2 capture process by absorption-regeneration into amine(s) based solvents applied to cement flue gases (the Brevik Cement Plant will be taken as reference): the first step will be to simulate the conventional process and the second step will be to adapt the model for the simulation of demixing solvents; - some experimental absorption-regeneration tests (lab/micro-pilot scale) with demixing solvents currently investigated at UMONS or any other experimental tests will be envisaged to have some inputs useful for the simulations; - technico-economic investigations in relation with the demixing solvents technology will be also necessary in order to optimize the application of such process in the cement industry. 2°) Note that from an administrative and official (UMONS) point of view, the second PhD thesis annual academic committees (“Comités d’Accompagnement de Thèse (CAT)) with internal (UMONS) and external members: → took place on 07-01-2016 for Nicolas Meunier and on 11-03-2016 for Sinda Laribi (these

committees have to meet after roughly one year of PhD thesis, then annually up to the end of the thesis);

→ were both validated. Both reports have been sent to ECRA.

1.1.2. Association of students works

One of the objectives of ECRA Chair is also to associate student works in the scientific activities of the Chair. In this context, besides the four PhD theses with large scientific contents, works of undergraduate students are also achieved, related to more specific subjects as parts of the scientific topics included in the ECRA Chair. Two Master Theses are currently under progress: → « Modeling and optimization of PSA processes for the treatment of gaseous effluents rich in CO2 » Nicolas DEBAISIEUX (student from UMONS) - Master thesis (February - June 2016) achieved in the Thermodynamics Unit; the student is working partially with N. Meunier. The report written in French will be available in June 2016. → « Technical, economical and environmental evaluations of CO2 capture techniques » Lucas LE MARTELOT (student from ESCOM, Compiègne, France) - Master thesis (February - August 2016) achieved both in the Chemical and Biochemical Process Engineering and Thermodynamics Units (3 months in each unit); the student is working alternately with S. Laribi and R. Chauvy. The report written in English will be available in August 2016.

1.1.3. Post-doc works

As described in the previous reports, in addition to the works linked directly to the ECRA Chair Scientific Coordination and logistic support (scientific support to the PhD theses, support for


reporting and publications writing, logistic organization of the Chair for event organization and support to any other scientific activities related to the Chair topics), Lionel Dubois is carrying out different post-doctoral works. The subject of these works are more specifically linked to HeidelbergCement activities. Note that HeidelbergCement is providing a complementary funding to the ECRA Chair (significant funding of the post-doc from January 2014 to December 2015 and co-funding of the fourth ECRA Chair PhD Thesis since January 2016). During the last year (from May 2015 to April 2016), the specific tasks carried out by the post-doctoral researcher in the framework of the HeidelbergCement collaboration were:

(1) a complementary bibliographic study on CO2 mineralization process into olivine derived compounds, especially on two aspects: the pH-swing mineral carbonation process and the utilization of cement wastes as feedstock for CO2 mineral carbonation.

(2) the continuation of the simulation works with Aspen Hysys software on the post-combustion CO2 capture process applied to Norcem Brevik cement plant flue gas: the use of alternative process configurations was investigated for MEA 30 wt.% (the investigations relative to PZ 40 wt.% and MDEA + PZ blend are still under progress).

Regarding the study on membranes CO2 capture process (investigated by the post-doc during the last years), it was decided to only carry out a regular technological monitoring on hybrid CO2 capture system combining membranes and amine absorption technologies but no specific tasks were assigned on this topic.

A summary of the more relevant information in relation with the post-doc tasks carried out during the last twelve months is given in Chapter 6.

1.2. Prolongation of the ECRA Chair at UMONS

In order to ensure a smooth continuation of the existing theses, to enable additional PhD theses to be assigned by UMONS, to complete the whole topics of the Chair and to keep the post-doctoral position for the Scientific Coordination of the Chair, a prolongation of the ECRA Chair was required and suggested by UMONS. The “contract” implies two consecutive periods of three years: from 2013 to 2016 for Phase 1 and from 2016 to 2019 for Phase 2 (see Figure 3). The final approval of ECRA’s Technical Advisory Board fell in May 2015, with a complementary funding by HeidelbergCement, allowing the current situation with four PhD theses and one post-doc.


Figure 3: ECRA Academic Chair timeline

1.3. Scientific Committee meetings ECRA-UMONS

As stated in the convention, two meetings of the scientific committee have to be hold during the year. Minutes are systematically written to summarize the discussions and decisions.

1.3.1. Meeting (ECRA – UMONS) at Mons on 22th October 2015

A meeting relative to the following aspects: 1. Global status of the ECRA Academic Chair activities (D. Thomas)

2. Intermediate reports on the ongoing PhD theses and post-doc (S. Laribi, N. Meunier, R.

Chauvy and L. Dubois)

3. Status of the recruitment procedure for the ECRA Chair PhD-4 (L. Dubois)

4. Signing of the ECRA Academic Chair prolongation for the 2016-2019 period (M. Schneider,

D. Gauthier, C. Conti and D. Thomas)

took place at Mons on 22th October 2015 (see Figure 4).

Figure 4: Meeting ECRA - UMONS at Mons on 22th October 2015


The dates (9 and 10th November 2016) and the global schedule (two-day event, in Mons and Lixhe) of the second ECRA Chair Event were also discussed.

1.3.2. Meeting (ECRA – UMONS) at Lixhe on 13th April 2016

Aiming at evaluating the progress of the different research projects of the Chair and discussing the research works, a meeting of the Scientific Committee took place at Lixhe on 13th April 2016. Here is the program of this meeting:

1. Report on ECRA’s activities beyond the ECRA Chair (M. Schneider)

2. Global status of the ECRA Academic Chair activities (D. Thomas)

3. Presentation of the new ECRA Chair PhD student (S. Mouhoubi) and intermediate reports on the ongoing PhD theses and post-doc (S. Mouhoubi, S. Laribi, L. Dubois, N. Meunier and R. Chauvy) 4. Discussions on the next ECRA Chair Scientific Event at UMONS (Global discussion) 5. Presentation of the Lixhe Cement plant 6. Visit of the plant (J. Wart)

Figure 5: Meeting ECRA - UMONS at Lixhe on 13th April 2016

1.4. ECRA Prize

An “ECRA Prize” was created to be awarded to new graduated students for the best Master2 project/Master thesis related to the CO2 capture or reuse, or related to any improvement for the cement industry. The second ECRA Prize (Prize amount: 400 EUR) has been awarded in September 2015 at the Polytech Mons Day (see Figure 6) to a deserving student: Guillaume Pierrot - “Experimental and simulation study of CO2 absorption into amine(s) based solvents : application to cement flue gases coming from partial oxy-fuel kilns ” – best Master2 project 2014-2015 This master thesis was achieved in the Chemical and Biochemical Process Engineering Unit.


Figure 6: ECRA award for a new graduated student (G. Pierrot) in September 2015

1.5. ECRA Chair website

As discussed during the meeting of 22th October, the website of the ECRA Chair was completed (see Figure 7), especially with a section including all the documents (works of students, annual reports… even if in French) of the ECRA Chair. The website’s address is: http://hosting.umons.ac.be/html/ecrachair/ and a password (available for all the members of the Scientific Committee) is needed to access the “Documents section”.

Figure 7: Illustration of the ECRA Academic Chair website

http://hosting.umons.ac.be/html/ecrachair/

ECRA CHAIR – Third Annual Report (April 2015 – May 2016) 15

2. ADVANCES IN THE PHD THESIS OF NICOLAS MEUNIER

This chapter will be divided into two different parts:

- The advances regarding the Dehydration Unit as a part of the Purification Process

applied to oxyfuel cement kiln flue gases;

- The advances regarding the Methanol Conversion Process as a competitive

opportunity for the reuse of CO2.

2.1. Advances in the Dehydration Unit

In the previous ECRA CHAIR Report [ECRA CHAIR Report, 2015], a dehydration unit was

presented as a part of the purification line for the oxyfuel cement kiln flue gases, and based

on a dual-bed Temperature Swing Adsorption (TSA) process.

Decisive improvements of our theoretic models have been performed so far including the

accurate description of the thermal and diffusive phenomena related to the adsorption of

water on the investigated adsorbents (i.e. zeolites 5A, 13X and silica gel). A particular attention

was also set to the creation of a reproducible methodology for the calculation of mass and

heat transfer coefficients. A master thesis is related to this subject (N. Debaisieux).

The next step regarding the study of the dehydration unit is to validate the theoretical results

obtained from our simulations with experimental data such as isotherm and breakthrough

curves for the different investigated adsorbents, especially regarding the particular behavior

of water on them.

Finally, as the study of the Sour Compression Unit is being finalized, the next report should

propose the overall Oxyfuel Purification Process with first approximations of key

parameters and possible ways of optimization.

2.2. Advances in the Methanol Conversion Process

This part will present the main advances regarding the conversion of CO2 into methanol and

will be divided into four topics:

- the description of the updated flowsheet of the methanol conversion unit;

- the discussion of the first simulated results based on the described flowsheet and

presenting the key parameters and consumption figures of the process;

- the first approximation of the operative (OPEX) and capital (CAPEX) costs generated by

such a process and the possible improvements currently investigated;

- the description of the micro-pilot plant being built in our laboratory to validate the

simulation results related to the kinetics of the chemical reactions occurring during the

methanol conversion process.

2.2.1. Description of the process flowsheet

In the previous ECRA CHAIR Report [ECRA CHAIR Report, 2015], a process converting CO2 into

methanol was presented, based on two catalytic reactors using a copper-based catalyst.


In this study, the CO2 conversion unit was simulated considering the conversion of a CO2-rich flow whose composition could be the one of a purified CO2-stream coming from the purification process applied to an oxyfuel cement plant. This composition is presented in Table 1, taking into account the hydrogen required for the conversion (stoichiometric H2/CO2 molar ratio equals 3). The flow rate considered in this study was also set to treat a CO2 partial flow rate of 866 tpd CO2 which represents 35% of the CO2 emissions of a typical cement plant producing 3000 tpd clinker [Corcoran, 2013]. As a reminder, the value of the CO2 capture rate from the cement plant flue gases was set to 35% because, regarding this CO2 capture rate, the production of the hydrogen required for the conversion of this amount of CO2 into methanol, would also produce enough oxygen to perform a complete oxyfuel combustion in the kiln.

Table 1: Characteristics of the gas entering the CO2 conversion unit

Gas composition (mol.%)

H2 74.59

CO2 24.87

N2 0.54

Temperature (°C) 30

Pressure (bar) 50

Flow rate (m³/h) 1677

Flow rate (Nm³/h) 75 560

Flow rate (kmol/h) 3298

In this process, the inlet synthesis gas is compressed (COMP-1) and preheated (EX-1) before entering the first reactor (REA-1) at the temperature and pressure of 250°C and 80 bar respectively. The reactor is adiabatic and consists of 600 tubes of 12.2 m length and 0.037 m diameter filled with 10 800 kg of a commercial copper-containing catalyst for methanol conversion. There is no recirculation in this first reactor. The outlet of the reactor is then cooled to preheat the inlet synthesis gas in the heat exchanger (EX-1) and fed to the second reactor (REA-2) which consists of 2600 tubes of 12.2 m length and 0.037 m diameter filled with 47 000 kg of the same catalyst. This second reactor is isotherm and the temperature is maintained at 260°C by circulating pressurized boiling water. The outlet of the second reactor is then cooled to preheat the inlet of the second reactor in a heat exchanger (EX-2) and further cooled (COOL-1) at the temperature of 40°C before being flashed (FLASH-1) to separate the water-methanol heavy mixture from other light gaseous compounds which are recompressed (COMP-2) and recirculated to the second reactor. A small fraction (2%) is also vented off to prevent the accumulation of inert gases (such as nitrogen) in the system. The liquid from the separator (FLASH-1) is pumped into a 25-stage distillation column on stage 15. The column operates at 1 bar and a reflux-drum temperature of 50°C is used so that cooling water can be used in the condenser. Two specifications are also set to the column: the purity related to the bottoms (0.01 mol.% methanol) and to the distillate (99 mol.% methanol). Finally, the small vapour fraction stream from the top of the reflux drum is compressed (COMP-4) and recycled to the inlet of the second reactor. Water and methanol streams are produced at bottoms and distillate respectively. A detailed flow sheet of the conversion process is provided in Figure 8.


Figure 8: Flowsheet of the methanol unit

2.2.2. Results and discussion

All the simulations have been performed with Aspen Plus v8.6 using the UNIFAC package for the calculations of gas thermodynamic properties. Considering the composition presented in Table 1 for the inlet gas, the resulting characteristics of effluents from reactors, compressors, and distillation column are presented in Table 2:

Table 2: Results from Aspen Plus simulations of the CO2 conversion unit

Composition (mol.%)

Reactor 1 outlet

Reactor 2 inlet

Reactor 2 outlet

COMP-2 outlet

COMP-4 outlet

CH3OH Product

Gas

H2 64.6 67.5 60.7 68.8 5.0 -

CO2 19.2 20.2 18.2 20.2 38.5 0.2

CH3OH 4.3 1.6 6.3 0.4 55.3 99.0

CO 3.5 3.5 3.1 3.5 0.4 0.8

N2 0.6 5.6 6.1 7.0 0.6 -

H2O 7.8 1.6 5.6 0.1 0.2 -

Temperature (°C) 285 240 260 48 190 50

Pressure (bar) 80 80 80 80 80 1

Flow rate (m³/h)

1804 8241 7778 4054 43 32

Flow rate (Nm³/h)

70 608 350 845 318 708 275 824 2030 -

Flow rate (kmol/h)

3039 15 109 13 847 11 947 124 767


Several observations can be pointed out from these results: 1) This installation is able to continuously convert 35% of the CO2 emissions coming from

a typical 3000 tpd clinker cement plant and produce a continuous flow of 32 m³/h (i.e. 584 tpd) methanol.

2) The presence of the flash (FLASH-1) is very important as it recycles more than 86 mol.% of the flow and prevent it from entering the distillation column

3) The presence of compressors 4 have to be analysed as it recycles only 124 kmol/h. Moreover, the compression of this effluent back to 80 bar is responsible for more than 75% of the electrical OPEX of the conversion unit (see section 2.2.3). Furthermore, the capital costs of compressors being one of the highest, the use of a flare to get rid of this small gaseous flow is currently under investigation.

Apart from the performances results presented in Table 2, the energy requirements for this conversion unit were also evaluated and key parameters and consumption figures are presented in Table 3. These results have to be considered very carefully as this process is neither optimized nor integrated yet. A particular attention will be also set on the reuse of the heat generated in the second reactor to decrease the thermal needs of the distillation column. Roughly, it can be estimated that this heat integration could reduce the reboiler heat duty by 19%. Moreover, the use of a flare could also reduce the electrical energy demand by 75% due to the huge amount of electrical power required for COMP-4.

Table 3: Energy requirements for the CO2 conversion unit

Parameters Value Unit

CH3OH production rate 24 ton/h

213 kton/year

Purge split ratio 2 %

H2:CO2 ratio 3 -

CH3OH yield (overall) 93 mol.%

Power of feed compressor (COMPR-1) 1.7 MW

Power of recycle compressor (COMPR-2) 0.9 MW

Power of recycle compressor (COMPR-4) 7.5 MW

Electricity usage 280 kWh/ton CO2

416 kWh/ton CH3OH

Electricity cost (at 0.08 €/kWh [Kiss et al., 2016]) 22 €/ton CO2

33 €/ton CH3OH

Heat generated (260°C) in the 2nd reactor (REA-2) - 9.1 MW

Heat duty of reboiler 17.6 MW

Steam usage for reboiler1 0.9 ton steam/ton CO2

1.3 ton steam/ton CH3OH

Steam cost (at 18 €/ton steam) 16 €/ton CO2

23 €/ton CH3OH

Energy2 generated in the cooler (COOL-1) (at 40°C) - 21 MW

1 Steam at 100 Psi (i.e. 6.89 bar) 2 Provided by free sources


2.2.3. Estimation of the operating and capital costs of the installation

A) Operating costs (OPEX): Operating costs are rather simple to estimate. Once the flows of the raw-material streams and the utility flows (fuel, steam, cooling water, power) are well-known, the operating costs are calculated by multiplying these flows by their respective euro (or dollar) value. However, care must be taken that the utility values are given on a thermodynamically consistent basis (i.e. fuel and electricity more expensive than high-pressure steam, itself more expensive than low-pressure steam, etc.). Otherwise, aberrations in prices could occur at times, so that it might appear that there is a profit in burning feedstocks to make electricity or in using electricity to produce steam [Douglas, 1988].

Considering the installation described in 2.2.1., the operative costs (OPEX) estimations are presented in Table 4.

It can be seen from this short operative costs analysis that the most expensive contribution

is caused by the reboiler heat duty which represents more than 73% of the OPEX (not considering the cooling water provided by free sources and the huge OPEX of COMP-4). As a result, heat integration will be a critical point in the optimization of the overall process in order to reduce this heat duty.

To this extend, Kiss et al. [Kiss et al., 2016] proposed an efficient process in which the heat generated in the second reactor could be reused to reduce the heat duty of the reboiler.

Table 4: OPEX estimation for the methanol conversion unit

Description Unit Rate Price (€/year)

Electricity

COMP-1 1.7 MW 1035

COMP-2 0.9 MW 575

COMP-4 7.5 MW 4600

Cooling water DISTIL-Condenser 14.5 Mton/y 4202

COOL-1 10.4 Mton/y 3002

Steam (at 100 Psi) DISTIL-Reboiler 268 kton/y 4400

This idea, among others, will be investigated in our future works on the optimization of the overall process. It also has to be mentioned that the cost of the catalyst is still not considered in the OPEX estimations as its turnover is currently difficult to estimate.

B) Capital costs (CAPEX):

There is a variety of ways of estimating the capital costs of equipment that range from very quick calculations with limited accuracy to very detailed calculations that are very time-consuming but more accurate. However, fast and simple approaches are required when envisaging conceptual designs and capital costs estimations are usually based on equipment cost correlations which are obtained by correlating a large number of provider’s quotes against the appropriate equipment size variable [Douglas, 1988]. Moreover, it has to be mentioned that equipment costs usually comprise two different prices:

- the purchased equipment cost, only considering the purchase costs of the

equipment;


- the installed equipment cost, considering the purchase and installation costs of the equipment.

B.1. Purchased equipment costs

A quite extensive set of correlations is available in Guthrie [Guthrie, 1969], Douglas [Douglas, 1988] or Peters and Timmerhaus [Peters et al., 1991] for the calculations of purchased equipment costs. These calculations often have the same expression (III.12.):

𝑃𝑢𝑟𝑐ℎ𝑎𝑠𝑒𝑑 𝐶𝑜𝑠𝑡 = (𝐵𝑎𝑠𝑒 𝐶𝑜𝑠𝑡)(𝐼𝑛𝑑𝑒𝑥)𝐹𝑐 (III.12.)

Where 𝐹𝑐 corresponds to a correction factor (e.g. for materials, pressure, etc.). The presence of an index is also required in these expressions to evaluate the cost increase of equipment with time. One of the most popular cost indices of this type is published by Marshall and Swift (M&S) and is updated monthly in the Chemical Engineering Journal [Douglas, 1988]. The Aspen® software also provides a more recent update for the calculation of base costs for a variety of pieces of equipment but also equations for their correction factors.

B.2. Installed equipment costs

Generally, the installed equipment costs is related to the purchased equipment costs by the addition of an installation factor 𝐼𝐹 (III.13.):

𝐼𝑛𝑠𝑡𝑎𝑙𝑙𝑒𝑑 𝐶𝑜𝑠𝑡 = (𝐵𝑎𝑠𝑒 𝐶𝑜𝑠𝑡)(𝐼𝑛𝑑𝑒𝑥)(𝐹𝑐 + 𝐼𝐹) (III.12.)

This installation factor is also tabulated according to the equipment considered in Guthrie [Guthrie, 1969], and other sources. Aspen® also has an updated version of these factors to provide a more accurate estimation of the installed equipment costs. Considering the installation described in 2.2.1., the purchased and installed costs (CAPEX) estimations have been calculated by Aspen Economics® and are presented in Table 5.

Table 5: CAPEX estimations for the methanol conversion unit

Unit type Unit Purchased costs (k€) Installed costs (k€)

Compressor

COMP-1 1439 1596

COMP-2 1221 1436

COMP-4 2568 2822

Reactor REA-1 278 576

REA-2 1316 1835

Heat exchanger EX-1 447 749

EX-2 2358 3612

Distillation column DISTIL 643 1324

Flash FLASH-1 70 171

Cooler COOL-1 174 312

Total 10 500 ~14 400


Once again, these values have to be considered with care and only as a first order of magnitude. As an example, further investigations are required to justify the presence of compressor COMP-4 considering the high CAPEX of compressors. A pie-chart is also proposed in Figure 9 to evaluate the repartition of the installed costs among the different unit types of the methanol conversion installation.

Figure 9: Repartition of the installed costs of the different unit type of the methanol

conversion installation

In this chart, it can be seen that the three most expensive parts of the process are compressors (41%), heat exchangers (30%), and reactors (17%).

B.3. Additional costs

Apart from the installed equipment costs, additional costs also have to be considered for the calculation of the total project capital cost. These additional costs include various items such

as piping, instrumentation, paint and escalation costs. Aspen Economics® also provide an estimation of these costs with regard to the simulated process. The additional costs to be considered for the simulated process are presented in Table 6:

Table 6: Additional costs for the installation of the methanol conversion unit

Total cost (k€)

Others 5035

Contingencies 4213

Piping 2705

Electrical 1135

Instrumentation 918

Remainings 1996

Total ~16 000

41%30%

17%9%

3%

Compressor Heat Exch. Reactors Distil. Column Others


Considering these additional costs, the total project capital costs are estimated to be 26 500 k€ which is a good approximation considering a 10 500 k€ total purchased equipment costs of the installation.

2.2.4. Description of the micro-pilot installation

In order to validate the results presented in the previous sections, a micro-pilot installation is being installed in our laboratory (Figure 10). This micro-installation will consist of the following units:

- three inputs streams (CO2, H2 and impurities (SOx, NOx)) with their respective flowmeters (D1, D2 and D3);

- a Booster Pump to increase the total pressure of gases up to 80 bar before entering the reactor;

- the catalytic reactor (with a fixed bed of catalyst) where the conversion of CO2 into methanol is occurring at 250°C and 80 bar

- a Back Pressure Controller used to keep the pressure constant along the installation; - a gas analysis device (Mass Spectrometer) to measure the concentration of the

different gases at the outlet of the reactor.

Figure 10: Schematic description of the micro-pilot installation

As presented in previous reports, operating parameters such as the pressure, temperature and Gas Hourly Space Velocity (GHSV) have major influences on the CO2 conversion yield. However, if temperature and pressure can be set and varied between experiments, the modification of the GHSV is more difficult as this parameter is directly linked to the dimensions of the reactor and the gas velocity. Moreover, as reducing the gas velocity could enhance diffusive resistances and make the determination of the intrinsic kinetic laws more difficult, the size of the reactor to be built has to be precisely studied. As a result, preliminary studies have been made to design and size the catalytic micro-reactor also considering the dimensions of the CuO/ZnO/Al2O3-type catalysts pellets used during the experiments. The key parameter in these parametric studies was to keep the GHSV at a value near 10 000 h-1 as it was proved to be the optimal working conditions for this CO2 conversion into methanol [Fournel et al., 2013]. Consequently, only two parameters can be varied keeping a pressure of 80 bar and a temperature of 250°C: the volume of the reactor and the feed gas flow rate.


Table 7 presents the relationship between the reactor volume and the feed gas flow rate at its inlet:

Table 7: Relationship between the reactor volume and the feed gas flow rate at constant GHSV, temperature and pressure

Diameter of reactor (inch)

Length of reactor (cm)

Feed flow rate (m³/h)

Gas velocity (m/s)

GHSV (h-1)

1

3 0.004 0.002

10 000

25 0.030 0.017

50 0.061 0.033

2

3 0.015 0.002

25 0.121 0.017

50 0.243 0.033

With these results, it can be seen that increasing the length of the reactor (and thus the volume of the reactor) will lead to an increase of the feed flow rate to keep the GHSV constant. Two other aspects also have to be considered: the minimum length of the reactor (with the minimum gas velocity) required to avoid the predominance of diffusive phenomena and the operative costs generated by high feed flow rates (especially considering the price of hydrogen). From these perspectives, it was set that the reactor should have a diameter of 1” and a length of 50 cm to reach an optimal GHSH of about 10 000 h-1, considering a feed flowrate of 0.061 m³/h (2.5 Nm³/h) and a stoichiometric H2/CO2 ratio (i.e. H2/CO2 = 3). Moreover, the reactor will also be built so that its length could be modified between experiments to vary the GHSV and thus enable the study of its influence on the CO2 conversion yield, keeping all other parameters constant. Figure 11 presents a schematic design of the reactor to be used in the micro-pilot installation. This reactor will consist of the following parts:

- two thermocouples to check the temperature during experiments; - two thermal loads to avoid temperature changes in the area of the inlet and outlet

of the reactor; - two porous materials to fix the thermal loads and allow the gas flow circulation; - the furnace/heated jacket to increase the temperature of the reactor to 250°C; - the reactor itself, filled with catalyst pellets, where the catalytic reaction occurs.


Figure 11: Schematic design of the micro-pilot reactor

Finally, the specifications for the experimental installation including flowmeters, back pressure controller, micro-reactor and heating device have been settled and sent to providers.

2.2.5. Conclusion and perspectives

Considering a CO2-rich flow whose composition could be the one of a purified CO2-stream coming from the purification process applied to an oxyfuel cement plant (but also from an absorption-regeneration process for CO2 concentration), it has been presented that the investigated process could achieve the continuous conversion of this gas to provide a CH3OH-rich stream with a purity of 99 mol.%. The energetic key parameters, operative and installed costs have also been presented as first estimations for the design of an integrated CO2-to-methanol conversion process. The next step for this project is to build the micro-scale experimental installation and to begin the experiments on different CuO/ZnO/Al2O3 catalysts and considering CO2/H2 mixtures (firstly without the presence of inert gas or gaseous impurity). During these experiments, some pertinent modifiers will also be investigated to evaluate their influence on the catalyst performances, stability and lifetime. The results obtained from these experiments will then be used to update the kinetic laws and allow a better understanding and simulation of this CO2 conversion process. In parallel, the optimization of the methanol conversion process will also be carried on in order to reduce its related operative and capital costs.


3. ADVANCES IN THE PHD THESIS OF SINDA LARIBI

In order to reach the purpose of CO2 conversion, several capture and purification techniques are studied in this work. Mainly, two fields are investigated:

post-combustion CO2 capture applied to an O2-enriched air combustion (partial oxyfuel combustion);

CO2 Purification Unit (CPU) applied to a full oxyfuel combustion. The diagram hereafter illustrates the two types of combustion cited, compared to the conventional combustion. The conventional combustion represented in black in Figure 12 is achieved with air, with a CO2 concentration in the outgoing flue gas between 20 and 30%. The partial oxyfuel combustion capture (oxyfuel combustion capture) is a combined technology working with a O2-enriched air combustion of the fuel, hence partial oxyfuel combustion, allowing a CO2 more concentrated flue gas (20 %<YCO2<70 %) compared to a conventional combustion. This type of combustion is represented in red in Figure 12.

Figure 12 : The types of combustion technologies considered

The full oxyfuel combustion capture system consists in realizing the combustion of the fuel using only oxygen, hence requiring an Air Separation Unit (ASU). The objective is to work with an outcoming flue gas from the cement plant concentrated in CO2 (YCO2= 70-90%). This type of combustion is represented in blue in Figure 12.

Figure 13 summarizes, for the two carbon capture technologies selected, the points deeply investigated in this PhD thesis.


Figure 13: Scheme of the two pathways detailed in this report (in red)

3.1. Partial oxyfuel combustion capture

3.1.1. Absorption-regeneration process

The absorption-generation process is a double step process represented in Figure 14: • The first step is the reactive gas-liquid absorption at typically 40°C: the flue gas coming from the cement industry is fed at the bottom of the absorber and flows to the top counter-currently with the solvent. It flows through the packing of the absorber, making contact with the solvent as it flows down. The CO2 is absorbed by the solvent as the flue gas rises so that the gas that comes out at the top of the tower contains very little CO2.

Figure 14 : Absorption-regeneration process

The second step is the CO2 desorption at 120°C for aqueous MEA 30 wt%: in the second

column, the temperature is increased by means of the energy produced in the boiler in order

to facilitate the CO2 desorption (the reaction between CO2 and amine being reversible). The

solvent drops in the desorber and the CO2 is released from solvent. Concentrated CO2 is

recovered at the top of the column and the lean solvent flows to the bottom of the desorber.


The energy demand for the CO2 regeneration has to be minimized through the sizing of

installations and the determination of operating conditions. The lean solvent is then recycled

to the absorber. The solvent flows from one column to another, with steady supplies of fresh

solvent to compensate the degradation of the absorption solution like thermal degradation,

oxidative degradation and degradation due to irreversible reactions with CO2.

3.1.1.1. Solvents for the gas-liquid absorption process

Different categories of solvents are selected for the screening of solvents under high CO2 concentrations, most of them are alkanolamines called in this report “amines” including primary alkanolamines, secondary alkanolamines and tertiary alkanolamines. Depending on the structure of the amine, they react differently with CO2, hence with different values of kinetic constants leading to different absorption performances. A simplified kinetic expression is used, with a kinetic constant (k) relative to each type of amine studied in this work: the CO2 reaction with an amine can be considered globally as an apparent 2nd order reaction with first orders relatively to each reactant CO2 and amine according to (Versteeg et al., 1996): r=k.CCO2 CAmine.

The used solvents during the absorption tests are a combination between the chemicals

described in Table 1 and Table 2 of Annex 1 with two categories of solvents standing out:

conventional solvents and activated solvents.

3.1.2. Experimental tests conducted at lab scale

Different types of tests were conducted in the context of a master thesis (Pierrot, 2015) using the cables-bundle contactor. The experimental device has been described in the previous report (ECRA Chair first annual report, 2014) together with the tests achievable:

A. Continuous tests

In these tests the liquid and the gas are continuously fed into the contactor. There is only one circulation of the liquid through the column. Some minutes of test are sufficient to reach a steady state for the CO2 absorption efficiency (A and GCO2, abs). These tests lead to the comparison of kinetic performances of different solvents.

Continuous tests without pre-loading (10% <YCO2,in < 60% ; αCO2=0):

The flue gas outcoming from a O2-enriched air combustion process can contain up to 60%vol of CO2. For this reason, tests for each solvent were conducted by varying YCO2,in from 10 to 60%vol at intervals of 10% in order to observe the influence of a high CO2 content on the absorption performances. The range of CO2 concentrations chosen is not limited to hybrid conditions. The tests at 10% allows obtaining the behavior of amines for a gas issued from power plant, the 20 and 30% points correspond to conventional cement plants conditions and the higher CO2 contents (40 to 60%) are relative to the hybrid conditions.

B. Semi-continuous tests (YCO2,in =40% and αCO2≠0)

In these tests the gas is continuously fed into the column but the liquid (without any feed) is recycled. The recirculation in the contactor of a constant volume of 1.3 l of solvent has been


achieved. The objective is to characterize the temporal evolution of YCO2,out , GCO2,abs and the CO2 loading (αCO2) of the solvents. If the test is achieved with a sufficient duration, an equilibrium can be obtained between gas and liquid phases, leading to CO2 absorption capacity. As intermediate value of the working range of continuous tests, considering the flue gas composition from partial oxyfuel combustion, YCO2,in was fixed to 40% for all the semi-continuous tests. A liquid sample is withdrawn every 15 minutes to observe the evolution of pH and the CO2 loadings via the TC and IC measurements. For each test, the absorption efficiency can be calculated. This is defined as the difference between the amounts of CO2 in the in- and outgoing gases, divided by the amount fed:

A (%)= Gin YCO2,in - Gout YCO2,out

Gin YCO2,in 100=

GCO2,abs

Gin YCO2,in 100

Gin = molar gas flow rate at the inlet of the contactor (mol/h) Gout = molar gas flow rate at the outlet of the contactor (mol/h) At the equilibrium, the CO2 loading representing the maximum quantity of CO2 that has been absorbed by a mole of amine is quantified:

αCO2 ( mol CO2mol amine⁄ )=

CCO2

Camine

But this quantity αCO2 can be determined at every time of the absorption test.

3.1.2.1. Summary of the important results

For semi-continuous tests, the CO2 molar flowrate decreases with time as the solution becomes increasingly loaded in CO2 and this can be linked to the increase of the loading of the solvents as long as recirculated in the column. The CO2 loadings of the solvents at the end of the tests are given in Table 8.

Table 8: CO2 loadings of the solvents at the end of the tests

Solvents αCO2 (mol CO2/mol amine) after 90 min

Sim

ple

Solv

en

ts

MEA 30% 0.455

MMEA 30% 0.456

DEA 30% 0.245

MDEA 30% 0.154

AMP 30% 0.553

AHPD 30% 0.476

PZ 10% 0.696

TETRA 30% 1.080

Act

ivat

ed

solv

en

ts

[MMEA 30%+PZ 5%] 0.481

[MMEA 30%+TETRA 5%] 0.515

[DEA 30%+PZ 5%] 0.446

[DEA 30%+TETRA 5%] 0.407

[AMP 30%+PZ 5%] 0.381

[AMP 30%+TETRA 5%] 0.396

The comparison of the CO2 loading given in Table 8 highlighted that even if PZ leads to a high CO2 loading, TETRA allows absorbing more CO2 than the other solvents tested as it presents the highest αCO2 at the end of the test.


Figure 15 summarizes the experimental screening of solvents with an evaluation of the absorption performances of a wide variety of solvents in high CO2 contents conditions. The comparison has been done with the conventional solvent MEA (30%).

Figure 15: Best solvents screened in high CO2 concentrations compared to MEA 30%

The best solvent at t=0 (αCO2=0) having the highest GCO2,abs was [MMEA 30% + PZ 5%] but after the CO2 loading (t=90 min), the best solvents become [AMP 30% + PZ 5%] and TETRA. In conclusion of this step, the various best solvents have been selected for further investigation and comparison, considering the flue gas composition from partial oxyfuel combustion.

3.1.3. Micro-pilot tests for the best solvents screened

In the context of an internship (Lucas Le Martelot) at UMONS, as a second step of the work, for the best solvents screened during the lab-scale tests, micro-pilots tests were conducted in order to evaluate their absorption performances under high CO2 content conditions. The absorption-regeneration process described previously is tested in the micro-pilot scale unit (see Figure 16) using as first solvent the reference solvent MEA 30% and considering different CO2 concentrations in the flue gas (YCO2, in= 20%, 30%, 40%, 50% and 60%). The first results of these tests will be presented in this report. But other solvents’ performances are also currently under investigation for comparison.

0

1

2

3

4

5

6

7

8

9

MEA 30% MMEA 30%+PZ 5% DEA 30% + PZ 5% AMP 30% + PZ 5% TETRA 30%

t=0 min t=30 min t=90 min

YCO2, in = 40%

αCO2

=0 αCO2

≠0

GC

O2,

ab

s (m

ol C

O2/h

)


Figure 16 : Micro-pilot unit for absorption-regeneration tests (with operating range)

Two types of analyses are conducted during the absorption test runs, namely liquid and gas phases analyses.

3.1.3.1. Liquid phase analyses

Liquid analyses are conducted by discontinuous manual samplings every 15 minutes. pH measurements (see Figure 17 (right)) and quantifications of the CO2 loading of amine solutions (see Figure 17 (left)) are determined using a TOC-VCSH Shimadzu analyzer. This analyzer measures the carbon content over time. Total Carbon (TC) includes all forms of carbon present in solution, both organic (TOC) and inorganic (IC). The CO2 concentration can be calculated using the difference between TC of fresh and loaded solutions as:

CCO2 ( mol CO2l⁄ )=

TC loaded solution (mg C/l) - TC unloaded solution

1000MM carbon

The CO2 loadings of the solutions (αCO2), could be deduced by the formula described before.

Figure 17: CO2 loading (left) and pH (right) of the Lean and Rich solutions - G= 1111 l/h; L= 19

l/h; MEA 30%; YCO2, in=60%

8

8.5

9

9.5

10

10.5

11

11.5

12

12.5

0 15 30 45 60 75 90 105 120

pH

Time (min)

Lean solution

Rich solution

0.0

0.1

0.2

0.3

0.4

0.5

0 15 30 45 60 75 90 105 120

αC

O2

(mo

l CO

2/m

ol M

EA)

Time (min)

Rich solutionLean solution


3.1.3.2. Gas phase analyses

During all the tests, the concentrations of the gas phase at the inlet and outlet of the absorption column and at the outlet of the stripper are measured continuously through an IR analyzer, respectively YCO2,in, YCO2,out and YCO2, regen. The absorption efficiency A (%) can then be calculated by means of the formula cited before. Results for gas phase analysis are presented in Figure 18. A steady-state is reached after more than 1 hour.

Figure 18: Evolution of the CO2 concentrations and the absorption efficiency

during the micro-pilot unit tests (G= 1111 l/h; L= 19 l/h; MEA 30%; YCO2, in=60%)

3.1.3.3. Evolution of the energy of regeneration Eregen

For the different experimental tests conducted with the reference solvent MEA 30% for different YCO2, in= 20, 30, 40, 50, and 60% at liquid flow rates giving, by simulation, an absorption ratio of 90%. The value 𝑄𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 was fixed (2000 W), actual 𝐺𝐶𝑂2,𝑟𝑒𝑔𝑒𝑛 was

deduced for each case and : can be calculated. Respective results (experimental and simulation) are given in Figure 19.

Figure 19: Evolution of the energy of regeneration of MEA 30%

0

5

10

15

20

25

0 20 40 60

E reg

en(M

J/kg

CO

2)

YCO2, in (%)

Sim Exp

𝐸𝑟𝑒𝑔𝑒𝑛(𝑘𝐽

𝑘𝑔𝐶𝑂2) =

𝑄𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 (

𝑘𝐽ℎ

)

𝐺𝐶𝑂2,𝑟𝑒𝑔𝑒𝑛 (

𝑘𝑔𝐶𝑂2ℎ

)


The energy of regeneration decreases when increasing the CO2 content in the flue gas; this conclusion is relevant and shows the advantage of working with high CO2 content conditions, characteristic of the partial oxyfuel combustion.

3.1.4. Estimation of the total energy required under partial oxyfuel conditions

3.1.4.1. Calculation of the energy required for the Air Separation Unit under

partial oxyfuel conditions

Thanks to our ECRA-UMONS collaboration through the ECRA Chair, calculation sheets

corresponding to a O2-enriched air combustion represented in Figure 12 were provided by

ECRA. These allow to calculate:

A. The composition of the flue gas outcoming from an O2-enriched air combustion

(identical composition of the recirculated and the extracted flue gases)

B. The O2 flow rate to be supplied by the ASU and the energy necessary for this

production.

These calculations include as a first and sufficient starting-point global balances of the major components in the clinker burning process. Deduced results for EASU (in the case of a fixed amount of CO2 recovered and a fixed O2 input demand into the kiln) are presented in Figure 20.

3.1.4.2. Simulation of the CO2 capture using the absorption-regeneration process:

estimation of the regeneration energy under partial oxyfuel conditions

Simulations of the absorption-regeneration CO2 capture process were conducted in Aspen Hysys V8.6 also in the context of a master thesis (Pierrot et al., 2015). Tests correspond to the design of a pilot unit used in the European projects (CASTOR/CESAR), and applied to the case of the Brevik cement plant taken as base case for flue gas compositions. Since all the design and operating parameters are available in literature (which is not the case with most of the other installations), the CASTOR/CESAR pilot unit was selected as our case study with definite operating conditions.

The pilot is sized to handle a flow of 5000 Nm³/h at the inlet of the treatment line resulting in a flow of Gin = 4000 m³/h at the inlet of the absorption column (after removal of a large portion of water, cooling and compression).

The absorption ratio is defined to 90 mol.% (90% of the molar flow rate of CO2 entering the absorption column is recovered at the outlet of the regeneration column).

L = 21 m³/h.

These two data lead to a fixed amount of CO2 absorbed (or regenerated GCO2,reg) of 1.5 tCO2/h.

The CO2 recovered purity is fixed at 98 mol.% (classical value).

Solvent selected: MEA 30%

Modelling characteristics used for the simulations with Aspen Hysys V8.6:

Acid gas package

Thermodynamic models: Peng-Robinson (gas phase) and e-NRTL (liquid phase)

Reactions sets included in the package (validated by literature).


Inlet compositions taken for the case studies:

Inlet compositions of Brevik’s cement plant .

Inlet compositions provided by ECRA from results of partial oxyfuel combustion simulations.

Intermediate inlet compositions defined by interpolation between the two previous cases regarding YCO2,in.

Figure 20 : Global energy required for partial oxycombustion process

Figure 20 shows the advantages of the partial oxyfuel combustion for Eregen: for the same amount of GCO2, regen = 1.5 t CO2/h we have a saving of 24% on the CO2 regeneration energy if the CO2 content is increased from 20 to 44 %. However, we have also to take into account the cost of O2 production in the air separation unit to realize an O2-enriched air combustion which needs more oxygen-enriched air to produce more concentrated flue gas in CO2 outlet of the kilns. Consequently, the overlap of the figure obtained for the O2 production (EASU) and the one of CO2 capture (Eregen) leads to the total energy (ETOT) required for a post-combustion capture technology applied to O2-enriched air combustion. These “tools” have to be refined but will be very useful for further simulations. We can also infer that both experimental and simulation tests lead to a saving on the solvent

regeneration energy if the CO2 content of the gas is increased.

3.2. Full oxyfuel combustion capture

3.2.1. SOx/NOx reactions chemical mechanism

A complete chemical mechanism representing the SOx and NOx simultaneous absorption into water for pressurized systems, used to simulate reactions in both gas and liquid phases occurring in the Sour-Compression Unit, is detailed in Figure 21. It takes into account our last and very comprehensive bibliographic studies (Ajdari et al., 2015) that highlighted the pH influence on the importance of the interaction reactions pathways that may occur and the SOx/NOx interactions deriving complexes represented in Table 9.

Table 9: SOx/NOx complexes denominations and formulas

Complete denomination Raw formula Alias

Nitrososulfonic acid ONSO3− NSS

Hydroxylamine N,N-disulfonic acid HNO(SO3)22- HADS

Hydroxylamine N-sulfonic acid HONHSO3− HAMS

0

0.5

1

1.5

2

2.5

3

3.5

4

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9

E (G

J/tC

O2

)

YCO2 (% vol) in the flue gas outcoming from the kiln

EASU

Eregen

ETOT


Figure 21 : Complete SCU Chemical Mechanism for SOx and NOx absorption without (A) or

with (B) SOx/NOx interactions (reactions selected for 1<pH<4)

The chemical mechanism without interactions (A, in black) and with interactions (B, in green) were simulated to compare their absorption performances in pressurized systems.

3.2.2. Simulation results of the Sour-Compression Unit (SCU) absorption

performances

Figure 22 represents the simulated flowsheet of the Sour-Compression Unit simulated in Aspen Plus V8.6.

Figure 22: Aspen Plus flowsheet of the simulated SCU

The compositions of the outcoming flue gas from a cement industry working with an oxyfuel combustion given by ECRA calculations, taken as inputs in the SCU together with the design specifications used in the model for the De-SOx and De-NOx gas/liquid contactors based on ranges given by Air Products, are listed in Table 10:


Table 10: Input compositions and design specifications of the SCU contactors

Components Mole fractions Total gas mole flow (mol/h) 4764 (120 Nm³/h)

O2 (%) 3.27 Total liquid flow rate 10 l/min N2 (%) 11.11 L/G 0.082 Ar (%) 1.34 Column packing IMTP 25 mm H2O (%) 1.00 Column diameter (m) 0.15 CO (%) 0.04 Column packed height (m) 12 CO2 (%) 83.13

NO (ppm) 861

NO2 (ppm) 96

SO2 (ppm) 156

In this case and schematized in Figure 23 when considering the chemical mechanism without interactions (A), the SOx abatement rate in the first column is 33.39% compared to 99.39% in the case of considering the chemical mechanism with all the interactions (B). At the end, in both cases, the second column removes all the remaining SOx components resulting in a total removal of SOx for the global process. For the NOx components, the differences between the chemical mechanism without interactions (A) and the chemical mechanism with all the interactions (B) are slighter than the case of SOx abatement (see Figure 23). Similarly to the SOx components, in both cases, the second column removes all the lasting NOx components resulting in a total removal of NOx within the global process.

Figure 23 : Abatement rate of the SO2 and NOx components in the first column of the SCU for

mechanism (A) and (B)

The pH of the system calculated by Aspen PlusTM is between 1 and 4, which justifies the reactions choice made to build the chemical mechanism illustrated in Figure 21. Here again a model was developed allowing to estimate performances of the SCU process for different operating conditions, a modified design and leading to the essential energy consumption of the process.

33.39

96.4599.39 95.91

0

20

40

60

80

100

SO2 NOX

Ab

ate

me

nt

rate

fir

st c

olu

mn

(%

)

A B


4. ADVANCES IN THE PHD THESIS OF REMI CHAUVY

The purpose of this work is to evaluate, in parallel of the CO2 conversion into methanol already considered within the ECRA Chair, the potential of other options for the re-use of CO2 captured from cement plants. Indeed, a methodological selection has to be conducted in order to select at least three interesting CO2 conversion routes based on various criteria that have to be previously defined. These selected CO2 conversion routes will be then simulated using the adequate software Aspen PlusTM in order to obtain relevant technico-economical indicators and parameters for completing an environmental assessment, using the Life Cycle Analysis (LCA).

The methodological framework proposed for the 4-year project is presented in Figure 24.

Figure 24: Methodological framework

Identification of the CO2 conversion routes

Literature review, data, properties etc.

Pre-selection Reduction of the panel of CO

2 conversion routes

Selection Selection based on technical, economic, environmental aspects as well as market considerations

Simulation of the process Use of ASPEN Plus to simulate the selected alternatives

Economic OPEX, CAPEX

Methodology development Based on criteria and weighted matrix

Environmental assessment Use of SimaPro to perform a LCA

Methodology development Based on pass/fail grade

Process design

Sustainable CO2-based process design

Impacts of the process Identification of the hot spots, i.e. origin of the impacts

Optimization Optimization between process design, environmental impacts and economics aspects. Integration

Are the selected routes validated?

NO

YES

NO

YES

Is the process optimized?


As a multitude of processes and chemical reactions is identified in the literature [Otto et al.,

2015]; [ENEA, 2014], a pre-selection methodology is proposed in order to reduce the panel of

options and obtain a shortlist. The pre-selected routes will be then studied more in details

using criteria specifically developed for this work and based on technical aspects, economical

aspects, energetic aspects, environmental aspects and market considerations. This second

stage is still under progress and will be further discussed.

4.1. Methodology and development for the pre-selection

Following a complete literature review, an initial analysis is roughly conducted based on a

pass/fail methodology presented hereafter.

4.1.1. Definition of the criteria

This initial assessment is based on the following considerations undertaken in a “decision tree”

(Figure 25):

The estimated time needed to reach commercial technological maturity. Only the

routes that have a timeframe to deployment smaller than ten years are selected.

The technological maturity of the processes. As the goal is to simulate the most

promising routes using ASPEN Plus, only the routes that are at least validated in an

engineering/pilot-scale (Technology Readiness Level, TRL ≥ 6) are selected.

The conversion of large volumes. There is a net unbalance in the direction of carbon

dioxide emissions compare to its utilization. The selected routes must convert

important volumes as the cement industry emits yearly over 2 billion tons of CO2

[GCCSI, 2013]. The question raised is whether large volumes can be converted into

(useful) chemicals and non-chemical applications. According to the Report of ECRA,

2013, a typical cement sized plant (BAT plant) would produce 3 000 tons per day of

clinker, the most important component used. Between 2 400 and 2 700 tons per day

of CO2 are released through clinker production (2 475 tons of CO2 generate for a BAT

cement plant, [ECRA, 2013]). Generally, up to 90% of the CO2 released is targeted to

be captured. Following this, the routes considered have to lead to products produced

in large quantities.

Economic and energetic aspects, as well as the consideration on the overall CO2 emissions are

taken into account in the in-depth analysis.


Figure 25: Decision tree flowchart to obtain the shortlist

The Technology Readiness Level (TRL) scale is therefore used to establish the pass/fail grade.

Therefore, the state of the development of emerging technologies is assessed based on the

nine-point numeric TRL scale, defined by the Horizon 2020 Program of the European Union

and the U.S. Department of Defense (DoD) [DOE, 2011].

4.1.2. Identification of the CO2 conversion routes and their level of maturity

Around 60 start-ups, 90 projects and 25 Research Centers have been referenced at different

level of maturity and performances. In addition, over 30 routes are identified in the report of

ENEA, 2014.

Regarding the chemical CO2 conversion, carbon dioxide is currently used at industrial scale for

the synthesis of urea, salicylic acid, polycarbonates and cyclic carbonates. Hydrogenation of

CO2 leads to various CO2-based compounds depending on the catalyst used for the reaction.

Renewable methanol through CO2 hydrogenation is at pilot scale, the “George Olah” plant in

Iceland (CRI plant) being close to commercial scale. Methane, generating from CO2 and

hydrogen (methanation synthesis) is also close to commercial scale as an industrial-scale

facility was constructed by Audi in Germany [ECRA, 2013]. Syngas is a mixture of interest as it

can be used as intermediate feedstock for a wide range of hydrocarbons and can be a source

of hydrogen and carbon monoxide, even if the cheapest way to produce it is from natural gas

[Araujo et al., 2014]. The direct synthesis of dimethyl carbonate is not currently practical at

industrial scale but an alternative using methanol as intermediate has been developed. Some

prototypes and pilots have been reported for the production of other chemicals of interest

such as formic acid, ethylene glycol and lactones. Moreover, researchers are currently working

on the development of other various CO2 conversion pathways and catalysts for the

CO2 conversion routes

YES

YES

YES

NO

NO

NO

Shortlist for in-depth analysis

Is the route expected to operate within the near/middle term time-

period? (< 10 years)

Is the route validated at least in engineering/pilot scale? (TRL ≥ 6)

Does the route have the potential to convert large volumes of CO2?


production of other CO2-based compounds, such as carboxylic acids (acetic acid), carbamates

(linear and cyclic), formaldehyde, isocyanates etc. In addition, around 123 chemical reactions

involving carbon dioxide from literature have been recently reviewed by Otto et al., 2015. Fine

chemicals, such as methylurethane, 3-oxo-pentanedioic acid and ethylurethane, identified by

Otto el al., 2015, are not be taken into account mainly due to their market size limitation.

Concerning the mineralization, an overview on CO2 mineral carbonation processes has been

undertaken by Dr. Lionel Dubois [ECRA, 2015]. Projects involving ex-situ mineral carbonation,

where the process is chemically carried out, are therefore referenced to evaluate their level

of maturity. According to the atom used for the mineral carbonation, sodium carbonate,

sodium bicarbonate, calcium carbonate and magnesium carbonate are produced.

Finally, regarding the bio-fixation processes, biological organisms are used to convert CO2 into

a variety of products. They include bio-fixation of CO2 from algae and bacteria. Jajesniak et al.,

2014, have reviewed various CO2-utilizing microorganisms, such as photosynthetic and

chemolithoautotrophic organisms, which seem to display ability in assimilating CO2 and

converting it into complex molecules. Zhang, 2015, have reviewed current demonstration

projects utilizing flue gas to grow algae.

Figure 26 presents therefore the TRL for the main CO2-based products for the 3 main routes

of carbon dioxide utilization, i.e. chemical, mineral and bio-based routes.

Figure 26: Technology Readiness Level (TRL) of the main CO2-based products

0

1

2

3

4

5

6

7

8

9

Chemical routes of CO2

utilization Mineralization Biological

processes


4.1.3. Size of the CO2 utilization

In order to evaluate the potential of CO2 reduction, in ton of CO2 per ton of product, the

CO2/product mass ratio is calculated. It represents the amount of CO2 which is necessary to

produce 1 ton of a product and is based on the stoichiometry of the chemical reaction.

Therefore, the greater the ratio is, the more CO2 can be incorporated as feedstock for the

production of 1 ton of product. It is worth noting that process efficiencies are not taken here

into account. This value defines therefore the maximum theoretical quantitative potential of

CO2 reduction; carbon dioxide used as a feedstock.

Figure 27 illustrates the specific CO2 fixation amounts as feedstock for the production of CO2-

based compounds. It varies between 0.085 and 2.74 tCO2/tproduct, respectively for the urea and

the methane.

Figure 27: Mass of CO2 as feedstock per mass of product (t CO2 / t product)

A BAT cement plant producing 3 000 tons of clinker per day generates therefore +/- 2 475 tons

of CO2 per day is taken as reference. Assuming that the cement plant operates 330 days a

year, it would produce 0.74 Mt per year of CO2 that can be converted, considering that 90%

of CO2 released is captured; all these reference figures are lumped in Table 11 [ECRA, 2013].

0 0.25 0.5 0.75 1 1.25 1.5 1.75 2 2.25 2.5 2.75 3

Urea

Dimethyl carbonate

Ethylene carbonate

Propylene carbonate

Polycarbonates

Formic acid

Acetic acid

Acrylic acid

Salicylic acid

Methanol

Dimethyl ether

Ethanol

Methane

Formaldehyde

Calcium carbonates

Magnesium carbonates

Sodium carbonates

Microalgae

Polyurethane

Specific mass of CO2

(t CO2 / t product)


Table 11: BAT cement plant specifications

Clinker production CO2 intensity CO2 emissions Assumption 3 000 tpd

0.825t/tclinker 2 475 tpd [ECRA, 2013]

0.99 Mt per year 0.817 Mt per year Operating 330 days a year

0.735 Mt per year 90% of CO2 captured (target)

At this point, the potential of CO2 reduction in MtCO2 per year, which is the multiplication of the specific CO2 fixation amounts per mass of product (Figure 27) by the worldwide production in Mt per year, represents also the CO2 utilization potential. The total CO2 utilization potential expresses therefore the overall CO2 demand of the process taking into account the potential market size. It is therefore possible to evaluate the “equivalent number of cement plants” required by dividing the total CO2 utilization potential by the CO2 emissions of one BAT cement plant that can be converted, i.e. 0.74 Mt (Table 11). This number expresses therefore the order of magnitude of the potential size of CO2 utilization. These results are presented in the following Table 12.

Table 12: Size of CO2 utilization

CO2 based compound Production (Mt/y) Potential of CO2 reduction

tonCO2/tonproduct MtCO2 /y Equivalent No. of

cement plants

Acetic acid 0.733

Acrylic acid 0.611

Calcium carbonates 113,9 ¹ (2013) 0.439 50,08 68

Carbamates 5,3 ¹ (2013)

Dimethyl carbonate 1,6 ³ (2014) 1.466 2,35 3

Dimethyl ether 11,4 ¹ (2013) 1.911 21,79 29

Ethanol 1.911

Ethylene carbonate 0,2 ¹ (2013) 0.499 0.0998 <1

Ethylene glycol 16,5 4 (2013)

Formaldehyde 21 ¹ (2013) 1.45 30,45 41

Formic acid 0,6 ¹ (2013) 0.956 0.57 <1

Inorganic carbonates 200 ¹ (2013)

Magnesium carbonates 20,5 7 (2011) 0.261 5,35 7

Methane 2.743

Methanol 50 ¹ (2013) 1.375 68,65 93

Microalgae 1.8 5

Polycarbonates 4 ¹ (2013) 0.5 2 2

Polyurethanes >8 ¹ (2013) 0.3 ~ 2,4 3

Propylene carbonate 0,2 ¹ (2013) 0.431 0.086 <1

Salicylic acid 0,17 6 (2011) 0.319 0,054 <1

Sodium carbonates >62 ² (2013) 0.415 25,73 35

Urea 155 ¹ (2013) 0.085 13,18 17

References: 1: [Aresta et al., 2013]; 2: [ENEA, 2014]; 3: [Front Research, 2015]; 4: [G.V.R., 2014]; 5: [Huang, Tan, 2014]; 6: [Otto et al., 2015]; 7: [USGS, 2011]


Table 12 indicates that, if the whole methanol production was substituted using CO2-based processes, the direct consumption of CO2 would be yearly about 69 MtCO2, without regard to the additional CO2 generation from power or heat emissions, or reagents production. This value is theoretical and represents only the maximum quantitative potential of CO2 reduction. In addition, the CO2 captured from around 93 cement plants and converted into methanol would substitute the whole methanol production, based on a BAT cement plant and a captured ratio of 90% (Table 11, Table 12). Either converting in mass more carbon dioxide into methanol or producing more methanol via other (conventional) processes would cause an overproduction resulting to price drop for example. Similarly, if the whole polycarbonates production was substituted, the direct CO2 consumption would be about 2 MtCO2. The CO2 captured from around 2 cement plants would be then used for its complete substitution (Table 11, Table 12). These figures are important to keep in mind while evaluating the potential size of carbon dioxide utilization. The global value chain needs to be taken into account. It is worth noting that an update on the yearly production amounts in Mtons would be required to fit better the current situation.

4.2. Evaluation for the pre-selection

Numerous CO2 conversion routes that use CO2 as feedstock have been then reviewed and evaluated through the pass/fail method. It has been decided that the CO2 conversion routes that use intermediates (reagents) which can be produced themselves from CO2 as feedstock will not be selected for the next stage. Table 13 presents therefore the shortlist of the routes of interest for further analysis that mostly fulfil the requirements of the three questions. They denote suitable options that have to be in-depth studied through technical, economic, energetic and environmental aspects as well as further market considerations for future development in the cement sector.

Table 13: Shortlist of the routes of interest

CO2-based compound CO2-conversion process Conventional production Calcium carbonates Mineralization (mineral carbonation) Extraction (mining)

Ethanol Microbial process * Hydration of ethylene * Fermentation

Formic acid Electrochemical reduction / electrolyze Synthesis from methyl formate

Methane Catalytic hydrogenation Upgrade of raw natural gas

Methanol Catalytic hydrogenation

Steam reforming of natural gas - Syngas production - Crude methanol production - Distillation and purification

Microalgae Biological process NA

Sodium carbonates Mineralization (mineral carbonation) * Solvay Process * Use of the mineral Trona

It is worth noting that some products of interest, such as polycarbonates, ethylene carbonates and propylene carbonates, have been rejected within this study because of their market size. Nevertheless, their productions have almost reached the commercial scale and involved


several actors. In addition, if the fossil-derived epoxides used as co-reactant are substituted by bio-based compounds, the production of polycarbonates would become highly relevant. In order to analyze these pre-selected CO2-conversion routes, the following information are required for each of them (if available): Raw materials, Main product(s), Side product(s), Reactions, Catalysts, Process steps, Energy requirements, Operating conditions. The chemical and physical properties of the substances related to the process, such as the solubility, the molar volume etc., will be determined through simulations using the software ASPEN Plus.

4.3. Methodological selection

In addition to the pre-selection of the CO2 conversion pathways, an in-depth analysis is conducted, based on various criteria and indicators. As this part is still under progress, only the criteria and indicators’ definitions are discussed hereafter. Rating scales for the methodological selection and the complete evaluation will be presented later in another report. Nevertheless, some conclusions can be drawn and are presented in the following.

4.3.1. Definition of the criteria and indicators

The short-listed CO2 technologies identified through the first stage is analyzed via the following criteria specially developed to select at least three routes that are the most suitable relating to technical aspects, economical aspects, energetic aspects, environmental aspects and market considerations. Criterion 1: Maturity of the process

Indicator 1-1: Technological maturity. It indicates the level of maturity of the different

pathways. The TRL scale is used to indicate the state of the development.

Indicator 1-2: Timeframe to deployment. It defines the estimated time needed to reach

commercial technological maturity. The industrial expansion of the technology is highly

depending on economics and regulations.

Indicator 1-3: Risk and uncertainty. It indicates how risks and uncertainties may impact the

industrial expansion and the dynamic of growth; includes the barriers and the behaviors

(public opinion and private companies) which might emerge, as well as the variability in

policies and the impacts of incentives. A SWOT matrix is used to evaluate therefore the

strengths, weaknesses, opportunities and threats involved.

Criterion 2: Economic potential

Indicator 2-1: Size of the market. It identifies how big the overall market is today and its

potential growth.

Indicator 2-2: Market competition with other technologies. It evaluates if the CO2 use is price-

competitive with alternative technology, products or processes, achieving the same outcome,

including low-carbon alternatives e.g. hybrid and electric vehicles; other green building

products.

Indicator 2-3: Economic viability. It indicates if it is economically feasible, innovative and

sustainable in terms of financial resources. A CO2 conversion process is economically feasible

if the unit production cost is inferior or equal to a reference case, usually the conventional way

of production, or if the process is forecasted as being close to profitability [Roh et al., 2016].


Indicator 2-4: Energy costs. The energy bill, i.e. the cost of the energy that has to be invested

to convert carbon dioxide into a valuable compound, is roughly evaluated. This can be compare

with the price of the CO2-based product.

Criterion 3: CO2 reduction potential

Indicator 3-1: Specific mass of CO2 as a feedstock. It represents the mass of CO2 which is

necessary to produce 1 ton of a product, as determined during the first stage.

Indicator 3-2: CO2 avoidance potential. It defines the maximum quantitative CO2 potential

abatement for the entire substitution compared to the conventional synthesis. It is roughly

estimated as the product of the specific mass of CO2 and the production volumes. Conversion

routes producing products in high volume are likely to have greater demand for CO2, so greater

CO2 avoidance potential. This indicator relies on the hypothesis that all the conditions, in terms

of technical conditions, economics and regulations, are met to favor the industrial expansion

of the technology, which is currently not always the case [Otto et al, 2015].

Criterion 4: Environmental, health and safety performance

Indicator 4-1: Environmental potential for the production of CO2-based compounds. It is

based on the stoichiometric approach. Optimal mass balance can be evaluated using

stoichiometric amounts. By multiplying the corresponding impact data.

Indicator 4-2: Health and safety considerations. It represents an important part when

alternatives are studied and must be an integrated element as it can cause human suffering

and incur high costs. As part of all managerial development processes, this indicator is

therefore difficult to evaluate in the context of this project. Whenever possible, health and

safety issues are then referenced for each alternative in order to compare them in terms of

hazardous substances utilization, etc. The controls and preventive and protective measures is

beyond the scope of this project.

Criterion 5: Energetic performance and operating conditions

Indicator 5-1: Energy requirements. It refers to the energy requirements (mainly electricity

and sunlight) for CO2 conversion and H2, if involved. Notice that H2 needs to be produced from

processes that avoid to co-produce CO2. The CO2-based routes are then evaluated by

comparison with their respective reference cases, i.e. their respective conventional way of

production.

Indicator 5-2: Operating conditions. It refers to the key parameters of the processes such as

the temperature, pressure and kinetics.

Indicator 5-3: Conditions of the CO2 feedstock. It indicates the conditions of the entering CO2,

mainly the level of purity and the CO2 concentration, which usually supply energy and may

cause for example additional CO2 emissions and resource depletion.

It is worth noting that carbon dioxide sequestration could be also interesting to be studied. The duration of the mobilization of the CO2 in the products formed is defined as the period of time when the carbon dioxide is effectively “sequestered”. It corresponds roughly to the product lifetime. Mineral carbonation leads to permanent storage of the CO2. However, most of the CO2-based products do not provide long lifetime storage as carbon dioxide is released through combustion (CO2-based fuels production) or product degradation (CO2-based plastics production) [Bennett et al., 2014].


4.3.2. First output and outlook

First results tend to highlight the high potential of three routes which mostly fulfil the requirements of the selection criteria:

- The methanol via the catalytic hydrogenation, work in progress of Nicolas Meunier;

- The formic acid via electrochemical reduction;

- Both the calcium and sodium carbonates via mineral carbonation.

These CO2-based routes may merit therefore further analysis. The interest of the methanol route via catalytic hydrogenation has already been demonstrated by the in-progress thesis of Nicolas Meunier. Besides, the use of CO2 in methanol production is almost implemented at commercial scale; in Iceland, a large-scale pilot plant has been built by Carbon Recycling International (CRI). Regarding the formic acid, it is worth noting that even though its market is relatively small (Table 12), this chemical represents currently a molecule of interest due to its relative added value so that it has been decided to include it for next stages. For example, the ECFORM Pilot from DNV uses this technology [DNV, 2011]. Finally, both the calcium and sodium carbonates via mineral carbonation represent suitable options for the conversion of large amounts of carbon dioxide (Table 12). Nonetheless, questions are raised concerning mainly their kinetics, generally slow under ambient temperatures, that have to be assessed. Two companies, Calera and Skyonic, have been identified respectively for the production of calcium carbonates and sodium carbonates. They make the state of development of the technologies reaching a TRL of 7 for both the products. For information, the company Skyonic is also involved in the production of sodium bicarbonate (baking soda). Next steps for the study will be to validate these selected routes and complete the simulation of the processes in order to obtain the technico-economical indicators and parameters for completing the environmental assessment, following the methodological framework presented in Figure 24. It has also to be mentioned that through the Master thesis of L. Le Martelot first environment assessments (for the absorption-regeneration process with MEA and the pressure swing adsorption using a zeolite) will be achieved (using a specific tool named SIMAPRO).


5. BEGINNING OF THE PHD THESIS OF SELOUA MOUHOUBI

5.1. Thesis context and content

Anthropogenic carbon dioxide emissions are considered to be at the origin of the greenhouse

effect which causes global climate change. In order to mitigate the climate change, solutions

have been proposed, among them the post-combustion CO2 capture technology. Initial

researches in this field have led to (MEA)-based process. However, the MEA process is known

to be very demanding in term of energy consumption for solvent regeneration. Efforts have

been made to develop advanced solvents and processes to improve energy performances of

post combustion CO2 capture. Biphasic solvents seem to be promising for post-combustion

CO2 capture by lowering the energy required for solvent regeneration. Indeed, biphasic

solvents present a phase separation at given temperature during CO2 loading and form two

phases: CO2 rich and CO2 lean. Only the rich phase will be sent to the stripper thus reducing

significantly and advantageously the flow rate which has to be regenerated.

This phase separation can occur either with or without solid precipitation. In this thesis, we

will focus on systems without solid precipitation since those presenting a precipitate are not

easy to handle at the industrial scale.

This thesis will include the following tasks:

- a detailed bibliographic review on different phase change systems specially on

demixing solvents. The goal of this part is to identify some liquid-liquid phase

separation solvents which seem to present good skills for CO2 capture;

- simulations using Aspen PlusTM or Aspen HysysTM of the post-combustion CO2 capture

process using demixing amines. The first step will be to simulate the conventional

process and the second will be to adapt the model for simulating a simple demixing

solvent. At the end, the best demixing solvent will be chosen in the optimized CO2

capture process (the Brevik cement plant flue gas will be taken as a reference);

- in parallel to simulation, experiments will be carried out. Firstly, they will allow to

highlight and analyze the demixing phenomenon and then to calculate some useful

inputs for the simulation. In addition to separate absorption-regeneration tests,

combined tests will be performed with the existing micro-pilot (an adaptation of this

pilot to demixing solvents will be required);

- a technico-economic investigation will be conducted to optimize demixing process

specifically for the cement industry application.

During this thesis several scientific exchanges will be necessary with internal and external

collaborators.

5.2. Liquid-liquid phase change systems

Demixing solvents exhibit a liquid-liquid phase separation for given temperature and/or CO2

loading conditions, this phenomenon allow important energy savings by regenerating only the

rich phase and thus reduces the energy required at the reboiler. The various chemical systems

involving demixing solvents (their structure will be defined further in the text) found in the

literature are listed in Table 14 and summarized here after.


Table 14: Amines systems characteristics compared to MEA benchmark

Solvent / Technology

Solvent composition

Main characteristics

Technology advancement

Thermomorphic

biphasic

solvents

MCA (1M) + DMCA (3M)

+AMP (1-1.5M)

- lower viscosity

- reduction of sensible heat

- low regeneration temperature

- no need of steam stripping

- lower critical solubility temperature

(LCST)

Laboratory scale

DMX solvents

Alkyl-multiamines

Use of blend of amines

- high capacity of CO2 absorption

- fast separation of the two phases

- low corrosion

- good stability toward degradation

- low striping energy

- lower critical solubility temperature

(LCST)

Mini-pilot

France

Self-

concentrating

solvents

Amine + alcohol

- high absorption rate and high CO2

loading

- low regeneration temperature - low corrosion and degradation

Laboratory scale

{MAPA+DEEA}

blend

DEEA (5M) + MAPA (2M)

- high initial absorption rate and

loading

- elevated desorber pressure

- regeneration temperature 88°C

Gløshaugen pilot

Norway

{BDA+DEEA}

blend

BDA (2M) + DEEA (4M)

- high cyclic capacity

- high CO2 loading

Laboratory scale

{TETA+DEEA}

blend

TETA (1.5M) + DEEA (3.6M)*

*approximated concentration

- faster reaction rate

- reduction of overall energy

requirement

- low viscosity and volatility

Laboratory scale

5.2.1. Thermomorphic lipophilic amine solvents

Lipophilic amines are constituted of hydrophilic (amino) and hydrophobic (alkyl) groups. They

exhibit a liquid-liquid phase separation upon heating, and form an organic phase which acts

as an extractive agent removing the amine from the aqueous phase. The reaction leads to

dissociation of the carbamate and bicarbonate species in the aqueous phase. [Zhang, 2013]

studied a blend of amines composed of N-methylcyclohexylamine (MCA) as absorption

activator and N,N-dimethylcyclohexylamine (DMCA) with good skills for regeneration.

However, this blend forms a heterogeneous solution in the absorber. In order to overcome

this weakness a small amount of solubilizer such as 2-Amino-2-methyl-1-propanol (AMP) is

added to the solvent mixture to increase the phase change temperature.


The chemical structures are given in Figure 28.

Figure 28: Chemical structures of MCA, DMCA, AMP

The biphasic system process includes two important sections, the absorber where CO2 is

absorbed into a lipophilic homogeneous solution at 40 °C and the regenerator operating at

80°C. To exploit the low heat value for solvent regeneration, in place of steam stripping,

studies on new desorption techniques such as nucleation, agitation, ultrasound and extraction

have been conducted [Zhang et al., 2012].

5.2.2. DMX Solvents

DMX solvents have been proposed by IFPEN. Usually, DMX solvents are alkyl-multiamines

which upon CO2 loading, depending on temperature lead to a phase separation [Dergal, 2013].

A simplified DMX process is presented in Figure 29. The decanter is preferably positioned after

the amine/amine heat exchanger and before the regenerator in order to make decantation

easier due to reduction of liquid viscosity associated with the increase of temperature.

Figure 29: Simplified process flow diagram of DMXTM

It was found as a result of an evaluation of the DMX process applied to industrial pilots [Raynal

et al., 2014], that the pilot tested could be of great interest for CO2 capture.

5.2.3. Self-concentrating solvent

Self-concentrating solvent for CO2 capture was patented by Hu Liang [Hu, 2005]. In this

process the solvent is composed of a mixture of an amine as activated component and an

alcohol as extractive agent. After absorbing CO2, the solution splits into two phases in the

decanter. Only the rich phase is sent to the stripper and the regeneration takes place at 95°C.


5.2.4. {MAPA+DEEA} blend

A blend of two amines, a tertiary amine DEEA (Diethylaminoethanol (5M)), and a diamine

MAPA (N-Methyl-1,3-Propanediamine (2M)) (see Figure 30) was studied by [Liebenthal et al.,

2013]. It was found that this mixture presents a phase separation upon CO2 loading, MAPA

and CO2 contained in the rich phase (heavy phase) and a DEEA rich phase but lean in CO2 (light

phase).

Figure 30: Chemical structure of DEEA and MAPA

The main products quantified in this system are primary/secondary carbamates,

dicarbamates, carbonates and bicarbonates. This blend of amines was studied by [Pinto et al.,

2014] in the Gløshaugen pilot plant. The process configuration is shown in Figure 31.

Figure 31: The Gløshaugen pilot plant

According to the authors the system shows good performances and was successfully implemented in CO2SIM (software designed for CO2 simulation developed by NTNU).

5.2.5. {BDA+DEEA} blend

Another blend of amines, namely BDA (1,4-Butanediamine (2M), see Figure 32) with DEEA (diethylaminoethanol (4M)), was used by [Xu et al., 2013]. This mixture can also form two phases after CO2 absorption. The phase separation was due to the fast reaction rate of BDA with CO2 since BDA is a primary amine with a faster reaction rate with CO2. Also for the limited solubility of DEEA in the reaction products of BDA with CO2.


Figure 32: Chemical structure of BDA

5.2.6. {TETA+DEEA} blend

[Ye et al., 2015] made a screening study of mono-amine and polyamine blended with regeneration promoters (DEEA, DMCA). Several immiscible liquid phases were found and evaluated at 30°C for absorption and 80°C for stripping. Among tested liquid systems, triethylenetetramine (TETA) and diethylaminoethanol (DEEA) blend (molar ratio of TETA to DEEA was kept at 3/7) seems to be promising with good performances.

5.3. Summary and work prospects

Phase change systems are considered to be a promising technology for CO2 capture. Since the use of these solvents leads to the formation of two phases, a CO2 rich phase and a CO2 lean phase, it allows to regenerate only the rich one by saving the regeneration energy. This phase separation depends on composition, temperature and CO2 loading. At the first sight, the demixing solvents likely to present good performances for CO2 capture are constituted of a blend of amines. This blend contains primary or secondary amines, which functions as an absorption accelerator and a tertiary amine, which provides the required CO2 loading capacity. However, there can be other simple amines with good properties. This fundamental screening is essential in my researches. For my first year of research, I will focus on the following tasks:

- continuation of the bibliographic works with full thermodynamic considerations and focusing on liquid-liquid systems in order to identify the appropriate system(s) to study;

- achievement of experiments (experimental devices have to be built or purchased), firstly to highlight the demixing phenomenon and secondly to deduce some useful data for the simulation;

- simulations using Aspen plusTM or Aspen HysysTM of the post-combustion CO2 capture conventional process; this model will be further adapted for the simulation of a simple demixing solvent.


6. ADVANCES IN THE POST-DOCTORAL WORKS OF LIONEL DUBOIS

6.1. Complementary studies on CO2 mineralization into olivine-derived compounds

6.1.1. pH-swing mineral carbonation process

First of all, it is interesting to remind the principle of the “pH-swing CO2 mineral carbonation process” using, for example, ammonium salts, illustrated on Figure 33. This process consists of five steps: 1/ In the first step, NH3 is used to capture CO2 from flue gas to produce NH4HCO3. 2/ In the mineral dissolution step, 1.4 M NH4HSO4 is used to extract Mg from serpentine ground to a particle size range 75–125 µm. 3/ The Mg-rich solution is then neutralized by adding NH4OH, after which impurities in the leaching solution are removed by adding NH4OH. 4/ The Mg-rich solution reacts with the product from the capture step NH4HCO3 to precipitate carbonates. Since the formation and stability of hydro-carbonates is temperature dependent, MgCO3.3H2O (nesquehonite) can be converted to 4 MgCO3.Mg(OH)2.4H2O (hydromagnesite) at temperatures above 70°C. Precipitation of hydromagnesite results in a solution mainly containing (NH4)2SO4. 5/ After water evaporation, the final step is the additive regeneration, with the decomposition of (NH4)2SO4 at ≈330°C, and producing NH3 for the capture step and NH4HSO4 for the dissolution step.

Figure 33: pH-swing mineral carbonation process with recyclable ammonium salts (adapted

from [Sanna et al., 2014])

In comparison with a typical capture process where CO2 is first absorbed by chemicals (e.g. NH3) and then desorbed (to recover the sorbent) and compressed for transportation (leading


to high energy consumption for stripping and compression), as carbonates are directly used in the proposed mineral carbonation, there is no need for desorption and compression of CO2. Another advantage of this pH swing process is that it is able to separate three different products: silica, magnesite and iron oxide. This process could also be integrated with the chilled ammonia CO2 capture process, which has been demonstrated to capture more than 90% of CO2. The main drawback of the aqueous pH swing ammonium based process is represented by the large amount of water that needs to be separated from the salts during the regeneration step. The amount of water to be evaporated is still too high and alternative separation methods need to be investigated in order to make this process economically feasible. The purpose of the present complementary study was to summarize bibliographic information in relation with the cost analysis of the pH-swing CO2 mineral carbonation process. All the details of this study are available in [Dubois, 2015a] and only the main relevant information are summarized here. This cost analysis was mainly based on the works of [Wang, 2011] at the University of Nottingham (UK). A simplified process scheme for sequestering 1 ton of CO2 thanks to Carbon Capture Storage by Mineralization (CCSM) is illustrated on Figure 34 where mass flows of each component are indicated in order to have an idea of the global mass balance of the process.

Figure 34: Illustration of the mass balance of the pH-swing mineral carbonation process [Wang, 2011]

The analysis of [Wang, 2011], based on the carbonation of 1 ton of CO2, highlights the importance of solid to liquid ratio (S/L) (50 g/l and 300 g/l studied) and gives as total energy consumption: 1088 (50g/l) to 400 kWh (300 g/l), namely 3.92 GJ/tCO2 to 1.44 GJ/tCO2, the MVR (Mechanical Vapor Recompression) evaporation step being the most energy intensive step (2.53 GJ/tCO2). More precisely, the study was divided in OPEX and CAPEX analyzes, and was based on four “optimization experiments” and the highest carbonation efficiency of 96 % was obtained when the molar ratio of Mg:NH4HCO3:NH3 was 1:1.5:2. Note that in these calculation


the electricity price was fixed at 0.03 $/kWh, Serpentine price was fixed at 5 $/t and transportation cost was fixed at 2.5 $ for 100 km. The total costs (OPEX+CAPEX) corresponding to the optimum case of the pH-swing CO2 mineral carbonation process are given on Figure 35.The total cost, excluding profit from product sale, estimated by [Wang, 2011] is 72 $/tCO2, the largest costs being associated with chemical costs (35%), followed by energy costs (22%) and feedstock costs (21%). The value of the products from the process is an important parameter for the costs, since the mass of products is 3.8 times larger than the mass of CO2 sequestered (212.5 ton products compared to 56 tCO2).

Figure 35: Repartition of the CAPEX costs of the pH-swing mineral carbonation process using

recyclable ammonium salts [Wang, 2011]

As indicated in Figure 36, [Wang, 2011] gave an estimation of the profit linked to the sale of products obtained from the process but it is highlighted that the products and market size are uncertain.

Figure 36: Total costs of the pH-swing mineral carbonation process using recyclable

ammonium salts including benefits from product sale [Wang, 2011]


Moreover, the current market size of these products, namely Product 1 (> 46.9 wt.% Si), Product 2 (> 60 wt.% Fe) and Magnesite, are relatively small compared to the large amounts produced by the process. Based on the figures presented in Figure 36, it should be possible to have a huge profit thanks to the sale of the products coming from the CO2 mineral carbonation process. Even if the benefits from the sale of Product 1 and Product 2 seem “realistic” (namely 20 and 33 $/tCO2), the benefit from the sale of Magnesite is very important (764 $/tCO2) and unrealistic. Nevertheless, if the sale of this third product is fixed at the same amount such as the two other ones (between 20 and 33$/tCO2), the financial balance is still interesting and competitive for the industry applying this process. Note that [Wang, 2011] compared the results of his study with other costs analyses presented in literature. It was shown that the costs vary in a large range depending on the process route considered (Direct or Indirect aqueous process) and on the feedstock used (Serpentine, Olivine or Wollastonite). Moreover, the study of [Wang, 2011] has the advantage to combine both capture and conversion costs (which does not seem to be the case for all the other references). Globally, on the 6 references indicated, 4 of them present a cost ≤ 91 $/tCO2, and two of them a cost lower than 75 $/tCO2 (69 €/tCO2). If we compare these costs with the cost reported for different types of classical CCS chains applied to power plants, the range communicated by [IPCC, 2005] is 25-75 $/tCO2. Based on this figures, the combination of the mineral carbonation process with recyclable ammonium salts with the CO2 capture process seems totally competitive with the classical CCS chain. Based on this bibliographic study, it can be concluded that even if the OPEX and CAPEX costs of the pH-swing CO2 mineral carbonation process are significant, two major aspects are important: there are still several opportunities to reduce these costs and such process allows the production of three different salable products that can be directly financially valorized. These aspects justify the interest of carrying out several studies on this topic, especially for the application in the cement industry (whose the higher CO2 content in the gas to treat could be an advantage in comparison with the power plants).

6.1.2. Cement wastes utilization in CO2 mineral carbonation processes

Utilizing cement wastes as feedstock in a CO2 mineral carbonation process is a possibility

which is evocated in several papers such as in [Pan et al., 2012]. Indeed, several types of

cement wastes are investigated by different authors, namely: Cement Kiln Dust (CKD), Cement

Bypass Dust (CBD), Construction and demolition waste, Cement/Concrete waste and Blended

hydraulic slag cement (BHC). Depending on the waste, the CaO content can be different and

influence the CO2 capture. Figure 37 shows the relationship between CaO content and the

actual CO2 capture capacity of various alkaline solid wastes in the literature.


Figure 37: Illustration of the relation between CaO content and CO2 capture capacity of

different solid wastes [Pan et al., 2012]

The CaO contents were significantly high in these wastes, e.g., steelmaking slag (~30–60% wt. CaO), residues from Air Pollution Control - APC (Ca content up to 35%) and bauxite (4.8% Ca), cement kiln dust (~34–50% wt. CaO), oil-shale waste (CaO content up to 50%), and fly ash (53% CaO) from municipal solid waste incinerators. Along with greenhouse gas emissions from the steel manufacturing industry, considerable amounts of alkaline solid residues such as fly-ash (FA) slag, ultra-fine (UF) slag, basic oxygen furnace (BOF) slag, and blended hydraulic slag cement (BHC) are generated, which are either used in various applications or ultimately landfilled. Without detailing all the reaction mechanism provided in [Pan et al., 2012], Figure 38 shows the mole balance and equilibrium conditions for carbonation of alkaline solid wastes.

Figure 38: Illustration of the mole balance and equilibrium conditions for carbonation of

alkaline solid wastes [Pan et al., 2012]


Which is important to note is the fact that the fraction (αi) of each carbon species present is dependent on the solution pH. At a low pH (~4), the production of H2CO3 dominates, at a mid pH (~8) HCO3− dominates, and at a high pH (~12) CO3

2− dominates. Therefore, accelerated carbonation with solid wastes is favored at a basic pH due to the availability of carbonate ions. [Iizuka et al., 2004] proposed a new CO2 sequestration process using carbonation of waste cement, which indicated that calcium extraction increased with pressure. They also found that the influence of S/L ratio was significant, with the calcium extraction rate about 50% with an S/L ratio of 0.29 wt% and 22% with the ratio of 2.9 wt%. As highlighted in [Pan et al., 2012], the carbonation of alkaline solid waste has been proved to be an effective way to capture CO2. Table 15 compares carbonation efficiency for different types of cement wastes operated under various conditions. In addition, summaries of factors required for effective accelerated carbonation are presented in Table 16.

Table 15: Carbonation conditions and efficiencies for different cement wastes (C.D. =

Carbonation Degree)

Table 16: Factors required for an effective carbonation

Without analyzing all the results presented in Table 15 it can be noted that the highest carbonation level was obtained for Cement Kiln Dust, utilizing a Direct Carbonation Method in Glass absorber column and leading to 290 gCO2/kgwaste.


Regarding the factors required for effective accelerated carbonation presented in Table 16 it is shown that multiple factors affect the conversion of accelerated carbonation, as indicated by numerous investigations listed by [Pan et al., 2012]. These factors must be further understood, since they determine the economic viability of the technology as well as help to identify the conditions that are most favorable for the reactions. Particle size and specific surface area are the most important factors affecting the dissolution kinetics of any kind of material. Since slag grinding is expected to be a fairly energy intensive process, it is important to find out how large a particle size can be used. Average grain particle sizes of less than 100–150 μm are already in the optimum range for efficient carbonation. A particle size preferably of ~100 μm, or no more than 500 μm, was found to be optimal based on the kinetics studies of calcium leaching. Note that the CO2 content of the incoming gas did not seem to have a significant effect on the degree of CaCO3 precipitation. As expected, the reaction time was inversely proportional to the CO2 content of the ingoing gas flow: the lower the gas CO2 content, the longer the duration of the carbonation. The other factors influencing the carbonation are the temperature, the pH value and the surface activities of the feedstock. Other information regarding the use of cement wastes were presented in the report of (Dubois, 2015a) as for example the paper of [Iizuka et al., 2004] proposing a new scenario of CO2 sequestration by using waste cement as a calcium source for the carbonate formation. Waste cement is a calcium-rich waste product containing calcium in the form of calcium

silicate hydrate (such as C3S; 3CaO･2SiO2 ･3H2O) and calcium hydroxide (Ca(OH)2). Thus, waste cement is considered as a potential candidate as a calcium source for carbonate formation for CO2 sequestration in terms of the CO2 sequestration capacity. In the paper of [Iizuka et al., 2004], the carbonate formation rates from a sample of waste cement were determined by laboratory-scale experiments. Based on the results, a model process (see Figure 39) was designed for the treatment of CO2 emitted from a 100-MW thermal power plant, and evaluated the feasibility of the proposed process in terms of energy consumption and cost.

Figure 39: Illustration of a CO2 sequestration process using concrete wastes from [Iizuka et

al., 2004]


Based on the laboratory extraction experiments, the energy consumption for the proposed process by [Iizuka et al., 2004] to sequester CO2 emitted from a 100 MW Thermal power plants was estimated to be 1543 kWh/tCO2, and the cost was estimated to be 83 $/tCO2. This value is quite high but is comparable with some estimations given for other CO2 capture-sequestration technologies. Globally, it seems still difficult for the moment to really evaluate the potential of using cement wastes as feedstock for CO2 mineral carbonation process but the available articles confirm that this is clearly a possibility and that carrying out further investigations on this possibility would be interesting and relevant. As the exact composition and characteristics of each waste can differ a lot depending on the plant, it would be interesting to carry out a parametric study on several factors (particle size, specific surface area, surface activity, etc.) which are known to influence the carbonation efficiency. It would be also interesting to study the utilization of solid wastes (especially cement ones) as complement to the other mineral sources and not as only feedstock (such as envisaged in the current papers).

6.2. Comparison of various configurations of the absorption-regeneration process for the

post-combustion CO2 capture applied to Brevik cement plant flue gases

6.2.1. Introduction

The key point for allowing a large deployment of the post-combustion CO2 capture technology applying amines based absorption-regeneration process is clearly its cost. In order to improve this process and to reduce its cost, three ways are possible: the development of new solvents (new solutions or new blends), implementing more efficient equipment (new gas-liquid contactors or new packings), and using new process configurations in order to take advantage of a better energy integration and to reduce its energy consumption. [Le Moullec et al., 2014] give more details regarding the technical description and the interest of other process configurations. The present work focuses on this third possibility and especially on the energy savings linked to the implementation of advanced process configurations for a flue gas issued from a cement plant. The flue gas coming from Norcem Brevik Cement plant was considered and the CO2 capture pilot simulated in Aspen HysysTM software was based on the installation used during the CASTOR/CESAR European Projects [Knudsen et al., 2009], all the design and operating parameters being available in order to validate the simulation. The configurations considered are described in the next section and the conventional monoethanolamine (MEA) 30 wt.% was selected as solvent in a first step. In a second step, other solvents such as piperazine (PZ) 40wt.% and a blend composed of PZ and methyldiethanolamine (MDEA) will be selected for future simulations. Note that such as in [Le Moullec et al., 2014], the interest of other process configurations is conventionally highlighted for power plants and that quantifying it for cement plants is innovative.

6.2.2. Simulation of different process configurations 6.2.2.1. Aspen HysysTM modeling parameters

The modeling was developed in Aspen HysysTM v.8.6 software using the acid gas package and the conventional “efficiency mode”. The first solvent selected for our simulations is


monoethanolamine - MEA (HO(CH2)2NH2) 30 wt.% and the reactions included in the acid gas package are given in Table 17.

Table 17: Reactions included in Aspen HysysTM acid gas package for MEA

HO(CH2)2H+NH2 +H2O ⟷ HO(CH2)2NH2 + H3O+

2 H2O ⟷ H3O+ + OH-

H2O + HCO3- ⟷ H3O+ + CO3

2-

CO2 + OH- → HCO3-

HCO3- → CO2 + OH-

HO(CH2)2NH2 + H2O + CO2 → HO(CH2)2NHCOO- + H3O+

HO(CH2)2NHCOO- + H3O+ → HO(CH2)2NH2 + H2O + CO2

The acid gas package developed by Aspen allows to simulate the removal of acid gases as CO2 and H2S. It includes the physicochemical properties of these acid gases, water, amines alone (such as MEA considered in the present case) and also several mixtures (such as aMDEA), a rate-based calculation model and a makeup unit operation to compensate for losses in water and amine in the system. The thermodynamic models used are e-NRTL for the liquids and the Peng-Robinson equation of state for the gaseous phase. The CASTOR/CESAR pilot unit was selected as case study because all the design and operating parameters are available in literature, which is not the case with most of the other installations. The pilot is sized to handle a flow of 5000 Nm³/h (at the inlet of the treatment line). This flow of 5000 Nm³/h at the inlet of the line results in a flow of 4000 m³/h at the inlet of the absorption column (after removal of a large portion of water, cooling and compression). Among the parameters imposed, it must be pointed out: the CO2 recovered purity, fixed at 98 mol.% (classical value); and the absorption ratio, equal to 90 mol.% (90% of the molar flow rate of CO2 entering the absorption column is recovered at the outlet of the regeneration column) The dimensions of the absorber and the stripper, and the operating conditions in each column are also set and compared to the MTU ones used at Brevik in Table 18. Note that the linear pressure drop per unit height in the absorber and the stripper are fixed to 0.5 kPa/m, and that the total packing height is divided in stages of 1 m high. The present simulation works have several interests for ECRA and HeidelbergCement:

- all the parameters related to the CASTOR/CESAR pilot are available for carrying realistic simulations which is not the case with the Mobile Testing Unit (MTU) of Aker;

- the CASTOR/CESAR pilot has a CO2 capture capacity 10 times higher than MTU of Aker;

- the objective is complementary to the Aker one: Aker applied a precise process (configuration and optimized solvent) and in the present case the simulation will highlight the interest of alternative configurations.


The flowsheet developed in Aspen HysysTM is illustrated on Figure 40 for the conventional process configuration.

Table 18: Comparison between MTU and CASTOR pilots parameters

Figure 40: Aspen HysysTM flow sheet for the conventional process configuration

The flue gas leaves the Brevik cement plant at 165°C and 100 kPa. It is compressed to 120 kPa and cooled down to 40°C before entering the flow sheet on Figure 40. The conditioned gas (“gas to treat”, ≈ 4000 m³/h, see composition in Table 20) is sent in the absorber where the CO2 is captured by an amine-based solvent, aqueous MEA 30 wt %. in this study (“lean solution in abs” in the flow sheet). The amount of CO2 absorbed into the column is calculated using the “rate-based model”. At the outlet of the column, the “Rich solution before preheat” is pumped and preheated to 110°C thanks to the internal heat exchanger. Then, into the regeneration column, the gas is stripped thanks to the heating power and the CO2 is recovered at the top of the regeneration column (“produced CO2” in the flow sheet). The regeneration occurs at 200 kPa and at around 120°C (corresponding to the boiling point of aqueous 30 wt.% MEA at such pressure level). The condenser cooling energy is automatically adjusted in order to satisfy the CO2 purity specification. The regenerated solvent (“lean solution before cooling”) which still contains some CO2 is pumped through the heat exchanger in order to be cooled down. To compensate possible losses in amine and water at the outlet of the absorber and the regenerator, a makeup unit is added in order to automatically adjust the total flow of liquid


while reaching the desired concentration of amine. After this make up unit, the lean solution is cooled down to 40°C before entering in the absorption column and beginning a new absorption-regeneration cycle.

6.2.2.2. Types of process configuration improvements considered

Three categories of process improvements are envisaged, namely: 1. Absorption enhancement: the purpose of this process modifications is to increase the CO2 loading at the absorber bottom or to reduce excessive driving force in the absorber section. In the present case, the “Rich Solvent Recycle” (RSR) configuration (see Figure 41) will be considered.

Figure 41: Aspen HysysTM flow sheet for the Rich Solvent Recycle (RSR) process configuration

The principle of this configuration is to recycle into the absorber a part of the rich solution coming from the bottom of this column. The rich solution going back to the column can also be cooled down in order to promote the CO2 absorption. In addition to the liquid flow rate (liquid to gas ratio, L/G) that must be optimized for all the configurations, the other parameters that must be specifically adjusted for RSR configuration in order to minimize the energy consumption of the system (especially the solvent regeneration energy) are: the fraction of the rich solution recirculated, the temperature of the solution before re-injection into the column and the level of reinjection into this column.

6. Exergetic or heat integration: the general idea of these process modifications is to perform heat integration between the different process streams in order to reduce the heat losses and the solvent regeneration energy. The process modification corresponding to this category and considered in the present case if the “Solvent Split Flow” (SSF) configuration (see Figure 42), also called “Rich Solvent Splitting”. With this configuration, a part of the rich solution coming from the absorber is directly sent at the top of the regeneration column without being preheated by the internal heat exchanger. This arrangement leads to a modification of the temperature profile into the stripper (it is more smoothed than with conventional configuration) and the heat recovered from hot lean solvent is maximized. Furthermore, thanks to the cold solution injected at the top of the stripper, the condenser cooling energy is reduced. The parameters that must specifically optimized for SSF configuration are: the fraction of the cold rich solution by-passing the internal heat exchanger and the injection levels of the solutions (cold rich solution and preheated rich solution) into the stripper.


Figure 42: Aspen HysysTM flow sheet for the Solvent Split Flow (SSF) process configuration

3. Heat pumps: the idea of this process modification category is to use a heat pump effect in order to increase the heat quality provided to the system. As described by [Le Moullec et al., 2014], this effect enables the valorization of heat available at a too low quality level or, more generally, when increasing the quality level is profitable. Two configurations corresponding to this category are considered in the present study, namely the “Lean Vapor Compression” (LVC) (see Figure 43) and the “Rich Vapor Compression” (RVC) (see Figure 44).

Figure 43: Aspen HysysTM flow sheet for the Lean Vapor Compression (LVC) configuration

Figure 44: Aspen HysysTM flow sheet for the Rich Vapor Compression (RVC) configuration


The principle of the LVC configuration is as follows: the lean solvent at the bottom of the stripper is flashed in order to produce a gaseous stream (mainly composed of water and carbon dioxide) which is compressed and fed back to the stripper. Such process modification reduces the reboiler steam demand and cools down the lean solvent going to the internal heat exchanger. Furthermore, the lean solvent exits the flash tank at lower temperature and heats the rich solvent in the heat exchanger also to a lower temperature, making the top of the stripper a bit colder which reduces the cooling requirement in the condenser. Based on operational experiences such as in the CASTOR/CESAR project [Knudsen et al., 2009], this modification is generally accompanied with an expansion or a modification of the internal heat exchanger in order to reduce the hot pinch of this exchanger to 5°C. Indeed, even if injecting the rich solvent into the stripper at a lower temperature is beneficial in terms of condenser cooling energy, it is counterbalanced by an unfavorable larger temperature difference between the inlet liquid temperature and the boiling temperature. Moreover, as the vapor coming from the compressor installed after the flash unit is very hot (which could lead to a hot spot inducing degradation problems into the bottom of the stripper), a supplementary heat exchanger can be installed in order to cool down this vapor (to 120°C) while giving a preheating complement to the rich solution before entering the regeneration column. Regarding the “Rich Vapor Compression” (RVC) configuration, the principle is quite similar as the LVC one even if it is the hot rich solvent which is flashed instead of the lean solution in order to produce a gaseous stream sent into the bottom of the stripper. In addition to the liquid flow rate, the important operating parameter in relation with LVC and RVC configurations is the flash pressure. This pressure depends on the stripper pressure but it has to be noted that a too low pressure (under atmospheric one) is not advised for economic reasons.

As a conclusion of this section, it must be noted that four different configurations are considered in the present work, namely RSR, SSF, LVC and RVC, corresponding to the three categories of process modifications. The focus is put on these configurations because it does not imply too much modifications of the conventional process and some of them (such as LVC) have already shown interesting results for the application to power plants which have to be confirmed for the application to cement plants.

6.2.3. Results and discussion

First of all, it must be noted that the simulation method was previously validated with the use of CASTOR/CESAR projects results (application to power plant) and also with the use of another study (other pilot design but same modeling method) concerning the application to a cement plant (see [Gervasi et al., 2014] for more details). Concerning the methodology used in this work, the objective was to minimize the solvent regeneration energy (Eregen [GJ/tCO2]) defined as:

𝐸𝑟𝑒𝑔𝑒𝑛 = Ф𝑏𝑜𝑖𝑙𝑒𝑟

𝐺𝐶𝑂2,𝑝𝑟𝑜𝑑𝑢𝑐𝑒𝑑

where Фboiler [GJ/h] is the heat duty provided at the bottom of the stripper and GCO2,produced

[tCO2/h] the amount of CO2 generated at the top of this column (outlet of the condenser). Note that no compression of the produced CO2 is considered in the present case because in


accordance with ECRA, the focus is put on CO2 valorization options for which the level of CO2 compression can be different depending on the CO2 conversion process considered. Moreover, as the different configurations imply the optimization of several operating parameters (see Table 19), systematic parametric studies were carried out in order to identify the realistic conditions leading to the minimum of Eregen.

Table 19: Parameters considered in the parametric study for each configuration

Base case RSR SSF LVC RVC

L/G L/G L/G L/G L/G

Injection level into

the stripper

Re-injected

fraction

Split fraction Flash pressure Flash pressure

Re-injection level

into the absorber

Injections level of

the cold solution

into the stripper

Re-injection

temperature into

the absorber

Injections level of

the preheated

solution into the

stripper

All the details of the parametric study are available in [Dubois, 2015b] and only a summary is

presented here.

6.2.3.1. Optimization of the liquid to gas flow rate ratio for the conventional configuration

(base case)

For the absorption-regeneration CO2 capture process, it is conventional to optimize the (L/G)vol. ratio in order to minimize the regeneration energy. The simulated results for different (L/G)vol. ratio (which means different liquid flow rates due to the fact the gaseous flow rate was kept constant) are presented on Figure 45.

Figure 45: Regeneration energy as a function of the (L/G)vol. ratio


First of all, it must be observed that the “parabolic tendency” of this graph is quite conventional for such process. The minimum regeneration energy was identified for a liquid flow rate of 22 m³/h ((L/G)vol. = 5.56 10-3) leading to Eregen = 3.36 GJ/tCO2. This value is in the range of conventional values measured for power plants (between 3 and 4 GJ/tCO2, see for example [Knudsen et al., 2009]). Nevertheless, it must be pointed out that 3.36 GJ/tCO2 is close to the minimum values conventionally measured for MEA 30 wt.% which could be justified by the fact that a higher CO2 content in the gas to treat (20 vol.% for the cement plant considered in the present case in comparison with the range 5-15 vol.% for power plants) is favorable to the absorption process.

6.2.3.2. Global results analysis for the base case

The detailed results corresponding to the base case optimal operating conditions are presented in Figure 46 and Table 20. The results corresponding to all the configurations will be compared in Table 21.

Figure 46: Temperature profiles into the absorber (left) and the stripper (right)

Regarding the temperature profiles into the absorber and stripper presented on Figure 46, first of all it must be specified that the temperature mentioned for stage 17 of the absorber corresponds to the temperature of the inlet liquid (40°C) and that the temperature for stage 0 of the stripper corresponds to the boiler temperature (121.8°C). The profiles are quite conventional for such operation units even if the maximum temperature reached into the absorber (around 85°C) is a little bit higher than other values (75-80°C) generally measured. This could be linked to the higher CO2 content of the gas to treat (20 mol.% for the cement plant considered) in comparison with power plants (5-15 mol.%) which induces a higher heat of reaction. Concerning the other simulation parameters (see Table 21), in addition to the regeneration energy (namely 3.36 GJ/tCO2) commented in previous section, it can be seen that even if it is not an issue for the absorption-regeneration process, the condenser cooling energy (Econdenser) is significant (-1.94 GJ/tCO2) and reducing it thanks to the use of alternative configurations


would be also benefic in practice (reduction of the water flow rate circulating into the condenser). Regarding the consumption of the liquid pumps (Epumps) equal to 1.57 10-2 GJ/tCO2, it corresponds to only ≈ 0.5 % of the regeneration energy and is thus not significant for the evaluation of the overall energy consumption of the process. Moreover, the CO2 loadings value for the rich (αCO2,rich) and the lean (αCO2,lean) solutions are equal respectively to 0.506 and 0.211 mol CO2/mol MEA, which is conventional for MEA 30 wt.% even if it must be noted that the value slightly higher than 0.5 is possible thanks to the CO2 content of the gas to treat (≈ 20 mol.%). Table 20 presents the gaseous compositions of the gas to treat, the gas treated and the produced CO2.

Table 20: Gaseous compositions for the base case ((L/G)vol. = 5.56 10-3) in mol. fraction

Component Gas to treat Gas treated Produced CO2

N2 6.47 10-1 6.29 10-1 1.54 10-4

CO2 2.04 10-1 2.09 10-2 9.80 10-1

H2O 6.21 10-2 2.64 10-1 1.95 10-2

O2 8.56 10-2 8.33 10-2 3.72 10-5

CO 1.33 10-3 1.30 10-3 4.48 10-7

SO2 1.11 10-4 5.35 10-5 2.98 10-4

NO 4.74 10-4 4.61 10-4 4.45 10-6

NO2 1.77 10-6 4.18 10-7 7.13 10-6

MEA - 3.77 10-4 7.65 10-11

As no specific reactions were added concerning the other gaseous species (SO2, NO, NO2, etc.), the decrease of their concentrations into the absorber can only be associated to the modification of their solubility into the liquid phase. Regarding the MEA, only very small quantities are present into the treated gas and the produced CO2. These results confirm the absorption ratio of 90% fixed as simulation parameter, and also the fact that the produced CO2 contains 98 mol.% of CO2, the rest being mainly composed of water. It has to be highlighted that the presented results were obtained for a rich solution injected at stage 9 into the stripper. Indeed, as presented in [Dubois, 2015b] , injecting at level 9 or level 10 (top of the column) gives similar results in terms of regeneration energy while the condenser cooling energy is a little bit lower when the rich solution is not injected too close to the condenser, namely at stage 9.

6.2.3.3. Global comparison of the energy savings for the different process configurations

A summary of the simulations results corresponding to the optimal operating conditions for each configuration is provided in Table 21 and in Figure 47. Among the different process configurations simulated with MEA 30 wt.% as solvent, it was shown that the heat pump modification (namely LVC and RVC process modifications) gives the best regeneration energy savings (almost 14%) leading to Eregen lower than 3 GJ/tCO2. Regarding the two other categories, namely absorption enhancement (RSR) and the exergetic


integration (SSF) smaller energy savings were measured (between 4 and 8 %) which can be justified by the fact that even with the conventional process configuration, the rich CO2 loading of the MEA 30% is close to its equilibrium value, thus the potential for increasing this loading is very limited (which is normally the purpose of RSR configuration for example). Concerning the SSF one, the major advantage is linked to the decrease of the condenser cooling energy (which is also clearly decreased thanks to the LVC configuration).

Table 21: Summary of the simulation results of the different configurations

Base case RSR SSF LVC RVC

Operating conditions (L/G)vol,opt (m³/m³) 5.56 10-3 6.06 10-3 5.81 10-3 5.30 10-3 7.33 10-3 Split fraction rich sol. (%) - 35 26 - - Re-injection sol. temp. (°C) - 40 - - - Re-injection abs. stage (N°) - 4 - - - Hot sol. stripper stage (N°) - - 7 - - Cold sol. stripper stage (N°) - - 10 - - Flash pressure (∆p) (kPa) - - - 100 100 αCO2,rich (mol/mol) 0.51 0.51 0.50 0.51 0.47 αCO2,lean (mol/mol) 0.21 0.24 0.22 0.20 0.25 Energy consumptions Epump (GJ/tCO2) 1.57 10-2 1.57 10-2 1.57 10-2 1.58 10-2 1.58 10-2 Econdenser (GJ/tCO2) -1.94 -1.89 -1.02 -0.91 -1.52 ELVC/RVC,compressor (GJ/tCO2) - - - 8.28 10-2 13.57 10-2 Eregen (GJ/tCO2) 3.36 3.07 3.22 2.91 2.95 Eregen savings /Base case (%) - 8.5 4.2 13.4 12.1

Figure 47: Optimum regeneration energy for each process configuration (left) and energy savings linked to the use of alternative configurations compared to the base case (right)

Finally, even if not negligible, the pumping and LVC/RVC compressor energy consumptions have an order of magnitude clearly lower than the regeneration energy.

6.2.4. Conclusions and perspectives

In the context of reducing the CO2 capture costs specifically for the cement industry, the present study focused on the Aspen HysysTM simulation of different configurations of the absorption-regeneration CO2 capture process using amine based solvents (MEA 30 wt.% in the present case) and applied to the flue gas coming from the Norcem Brevik Cement plant. The design of the CO2 capture plant considered for the simulation was based on the CASTOR/CESAR European Projects pilot. Four process modifications were investigated, namely RSR, SSF, LVC and RVC, in order to be representative of the three categories of process modifications. For each configuration, a systematic parametric study on operating parameters


((L/G)vol. ratio, split ratios, flash pressures, etc.) was carried out in order to identify the conditions minimizing the regeneration energy (taken as comparison factor for each configuration). Based on the simulations carried out, it was highlighted that the heat pump modifications LVC and RVC lead to the higher energy savings while also reducing significantly the condenser cooling energy. The energy savings linked to RSR and SSF modifications were lower. As perspectives, other solvents will be considered for future simulations, such as piperazine (PZ) alone or used as absorption activator for methyldiethanolamine (MDEA), because the energy savings linked to the use of alternative configurations strongly depend on the solvent properties. Finally, the energy savings of the different process configurations and obtained with the different solvents will be globally compared. The possibility of studying other configurations or the combination of two configurations will be also envisaged.


7. GENERAL CONCLUSIONS AND PERSPECTIVES

7.1. General conclusions of progress of the ECRA Chair scientific activities

Different subjects related to CO2 capture and purification but also to CO2 re-use in the cement industry are addressed in this report, as a summary of the different research activities at UMONS during this third year (April 2015 - May 2016) of the existence of the ECRA Chair. As discussed and validated by the Scientific Committee, different scientific works, namely: - Extensive bibliographic research on various topics; - Specific technological monitoring;

for all the PhD students and post-doc.

- First contact and uses with simulation tools for PhD students 3 and 4 - In-depth simulation works for PhD students 1 and 2 on: o Flue gas purification for oxy-fuel processes, in particular the simulation of the Air Products

process with a sour compression step (SCU), a dehydration step and a cryogenic step in the CPU;

o Post-combustion CO2 capture, in particular the simulation of the conventional CO2 capture process into MEA, including alternative configurations;

o CO2 conversion into methanol with technology description (including OPEX and CAPEX estimations), kinetic study for simulation implementation, allowing the design of the experimental set up;

- Experimental works namely : o First CO2 capture processes investigations achieved on the amine absorption-

regeneration process, related to the possible interest for cement industry to consider a combination of O2-enriched conditions and post-combustion capture process: tests performances on several amine-based solvents at high CO2 content (up to 60%) into the flue gas

- Methodological selection of CO2 conversion routes (PhD thesis 3)

were carried on and are still under progress. Official reports and scientific communications (oral and poster presentations, abstracts and papers) were regularly written. All the planned scientific subjects (PhD theses, post-doc and Master theses) were gradually, and as consistently as possible, introduced into the overall framework of the ECRA Chair (see Figure 2).


7.2. ECRA Chair scientific activities for the next year

The following tasks are planned for the next year (period May 2016 – May 2017):

Concerning the PhD Thesis of Nicolas Meunier:

The focus will be put on the development of the micro-scale experimental installation in order to begin the experiments with different CuO/ZnO/Al2O3 catalysts and considering CO2/H2 mixtures (without the presence of gaseous impurity in a first step). The experimental results will be used to update the kinetic laws and to have a better understanding and simulation of this CO2 conversion process. The global objective of these works will be to optimize the CO2 conversion process into methanol, especially for the reduction of the operating and capital costs.

Regarding the PhD Thesis of Sinda Laribi:

The future works will be divided in two main parts. Firstly, concerning the partial oxyfuel combustion process, the absorption-regeneration tests using the micro-pilot unit will be continued for the best solvents previously screened and Aspen HysysTM simulations will be performed in parallel in order to quantify the regeneration energy savings thanks to higher CO2 content in the gas to treat. An evaluation of the impact of partial oxyfuel combustion conditions on the global chain (from O2 production to CO2 conversion) will be also investigated in terms of global energy consumption. Secondly, the simulated SCU process in Aspen PlusTM

considering SOx-NOx interactions with be optimized considering the energetic, environmental and economic aspects through a parametric study (variations of operational levels of pressure, water flow rate, recycle flow rate and design parameters). The final objective will be to take advantage of the pressure dependence of the rate determining reactions and of higher solubilities in the SOx/NOx system, in order to reduce the process to a single column absorber that could be applied to cement industries.

Relating to the PhD Thesis of Remi Chauvy:

In addition to the CO2 conversion into methanol studied by Nicolas Meunier, further analyzes of two other routes will be performed, namely the CO2 conversion into formic acid via electrochemical reduction and the CO2 mineral carbonation into calcium and sodium carbonates. The next steps for the study will be to validate the selected routes and to complete the simulation of the processes in order to obtain the technico-economical indicators and parameters for completing the environmental assessment, following the methodological framework established. Life Cycle Assessment (LCA) of CO2 capture processes applied to cement flues gases will be also envisaged in order to evaluate the environmental impact of the global chain , from CO2 capture to final CO2-based products.

About the PhD thesis of Seloua Mouhoubi

During the first year of the PhD Thesis, the focus will be put on three main tasks: the continuation of the bibliographic study on liquid-liquid systems including thermodynamic considerations in order to identify the chemical systems that will be further investigated, CO2 absorption experiments at lab scale will be performed with the select systems in order to highlight the demixing phenomenon and to calculate some useful parameters for the


simulation, and finally first simulations of the post-combustion CO2 capture using Aspen PlusTM or Aspen HysysTM and considering conventional process will be initiated with the purpose of adapting it in the future for a demixing solvent.

Regarding the Post-Doc works of Lionel Dubois:

In addition to the support works for the ECRA Chair, the post-doctoral works during the next year will be mainly focused on the continuation of the Aspen HysysTM simulations of the post-combustion CO2 capture process applied to the Norcem Brevik Cement plant flue gases, especially considering different process configurations and other solvents than MEA (namely piperazine (PZ) and PZ-methyldiethanolamine (MDEA) blend). The energy savings of the different process configurations and obtained with the different solvents will be globally compared, and the possibility of studying other configurations will be also envisaged.

7.3. External communication

7.3.1. Publications

Different abstracts/conference papers/manuscripts were published during the period May 2015-April 2016 and are planned for a next future. Past conferences and attached publications 1°) TCCS-8 – Conference on CO2 Capture, Transport and Storage - Trondheim, Norway –

June 2015 → “ Simulation of a CO2 purification unit applied to flue gases coming from oxy-combustion

cement industries ” Sinda Laribi, Nicolas Meunier, Lionel Dubois, Guy De Weireld and Diane Thomas → “ Solvent screening for the post-combustion CO2 capture applied to flue gases coming

from conventional and partial oxy-fuel combustion cement kilns” Guillaume Pierrot, Sinda Laribi, Lionel Dubois and Diane Thomas The two abstracts were accepted, with communication by posters. 2°) EMChIE 2015 - European Meeting on Chemical Industry and Environment - Tarragona, Spain - June 2015 → « Simulations with different process configurations for the CO2 capture applied to cement

flue gases » Julien Gervasi, Lionel Dubois and Diane Thomas A mini-article was accepted and is published in the proceedings. A poster was presented, awarded by a Best Poster Award.


3°) ICCDU 2015 - International Conference on Carbon Dioxide Utilization – Singapore - 05-09/07/2015 → “ Innovative solvents for the post-combustion CO2 capture absorption-regeneration

process applied to cement plant flue gases ” Lionel Dubois, Nicolas Meunier, Sinda Laribi, Julien Gervasi, Guy De Weireld and Diane Thomas → “ CO2 capture and re-use from oxyfuel cement kilns: process simulation of the CO2

purification and catalytic conversion into methanol ” Nicolas Meunier, Sinda Laribi, Lionel Dubois, Diane Thomas and Guy De Weireld The two abstracts were accepted for posters presentations. The first subject was rewarded by a Best poster Award.

4°) PCCC3 - 3rd

Post Combustion Capture Conference - 08-11/09/2015 – Saskatchewan, Canada

→ “ Post-combustion CO2 capture: optimization of the absorption-regeneration process for

the application to cement flue gases “ Guillaume Pierrot, Julien Gervasi, Sinda Laribi, Lionel Dubois and Diane Thomas Poster presentation → “ Post-combustion CO2 capture: screening tests of solvents for the absorption- regeneration

process applied to cement flue gases with high CO2 contents”

Sinda Laribi, Guillaume Pierrot, Lionel Dubois and Diane Thomas Oral presentation Two extended abstracts were accepted. The attendance to the congress was completed by a visit of Boundary Dam CCS plant (Saskpower).

Figure 48: Lionel Dubois and Sinda Laribi taking part to PCCC-3 Conference and to the Boundary Dam CCS plant visit in Canada


5°) ECCE-10 – European Congress of Chemical Engineering - Nice, France – September 2015 → “ Post-combustion CO2 capture applied to cement plant flue gases: screening tests of

innovative solvents for the absorption-regeneration process “ Sinda Laribi, Lionel Dubois and Diane Thomas → “ Optimization of a sour-compression unit for CO2 purification applied to flue gases

coming from oxy-combustion cement industries ” Sinda Laribi, Nicolas Meunier, Lionel Dubois, Guy De Weireld and Diane Thomas Two abstracts were accepted and 2 posters were presented. 6°) ICEF 2015 – Innovation for Cool Earth Forum – Tokyo, Japan – 7- 8 October 2015 → “Perspectives for carbon capture et reuse in the cement industry in Europe “ Diane Thomas A short conference (to the Cement session) was presented on ECRA Chair activities. A visit of Seikei University was also organized. 7°) PhD Day JJC GEPROC-UGéPE 2015 – Saint-Quentin, France - 15 October 2015 → “Post-combustion CO2 capture: Screening tests of solvents for the absorption-

regeneration process applied to cement flue gases with high CO2 contents“

Sinda Laribi Sinda has presented an oral communication and was rewarded by the Best Oral communication Award. → “Dehydration of CO2 coming from oxyfuel cement kilns by Temperature Swing Adsorption

(TSA) process “ Nicolas Meunier Nicolas has presented a Poster. 8°) BSDS 2015 – Brussels Sustainable Development Summit - Brussels, Belgium - 19-20 October 2015 → “Global optimization of the CO2 capture and reuse applied in the cement industry“

Lionel Dubois, Sinda Laribi, Nicolas Meunier, Guy De Weireld and Diane Thomas Presentation of an oral communication and a poster

9°) UTCCS-3 - 3rd

University of Texas Conference on Carbon Capture & Storage - 17-19/02/2016 – Austin, Texas, USA


→ “ Study of the post-combustion CO2 capture applied to cement plant flue gases with high

CO2 contents ”

Sinda Laribi, Lionel Dubois, Guy De Weireld and Diane Thomas → “ Simulations of various configurations of the post-combustion CO2 capture process

applied to a cement plant flue gas: parametric study with MEA 30 wt.%” Lionel Dubois, Guy De Weireld and Diane Thomas Two abstracts were accepted, linked to two oral presentations (remote presentations from UMONS). 10°) SCOT Workshop Infrastructure Managers – 23/02/2016 – Brussels, Belgium → “ ECRA Academic Chair research activities in relation with CO2 Capture & Reuse ”

An oral presentation was achieved by Lionel Dubois & Nicolas Meunier

7.3.2. Future planned communications

1°) FOA 12 – Fundamentals Of Adsorption – Friedrichshafen, Germany – 29 May - 3 June 2016 → “Dehydration of CO2 coming from cement oxyfuel kilns by temperature swing adsorption (TSA) using zeolites 5A, 13X and silica gel “ Nicolas Meunier, Lionel Dubois, Diane Thomas and Guy De Weireld Submission of two abstracts 2°) ICCDU 2016 – International Conference on Carbon Dioxide Utilization - 11-15 september

2016 – Sheffield (UK) → “CO2 re-use from oxyfuel cement kilns: Optimization of the CO2 catalytic conversion into

methanol” N. Meunier, R. Chauvy, L. Dubois, D. Thomas, G. De Weireld → “CO2 utilization from cement plant flue gas: Selection of suitable CO2 conversion routes for

the cement sector “ R. Chauvy, N. Meunier, L. Dubois, D. Thomas, G. De Weireld Submission of two abstracts

3°) GHGT-13 congress – Lausanne (Switzerland) - November 2016 → “Simulations of various configurations of the post-combustion CO2 capture process

applied to a cement plant flue gas: parametric study with different solvents” L. Dubois

and D. Thomas


→ “Study of the post-combustion CO2 capture applied to conventional and partial oxy-fuel

cement plants” L. Dubois, S. Laribi, S. Mouhoubi, G. De Weireld and D. Thomas → “Optimization of the Sour Compression Unit (SCU) process for CO2 purification applied to

flue gases coming from oxy-combustion cement industries” S. Laribi, L. Dubois, G. De Weireld and D. Thomas Submission of three abstracts.

7.4. Global perspectives of the ECRA Chair

The First ECRA Chair Scientific Event has been organized at UMONS on the 26th November 2014. This can be proclaimed as a great achievement and success as 100 participants from around 20 countries were registered. This event was a great opportunity to attend very interesting presentations given by industrial and academic experts, also by PhD students, on the subject of CO2 capture and reuse in the cement industry.

The Second ECRA Chair Scientific Event is planned at Mons on 09th November 2016, in the same “spirit” as the first one, followed by a visit of the CBR (HeidelbergCement) Lixhe Cement plant on the 10th November 2016. The draft of the event program is given in Annex. Note that the proposed global title is “CO2 capture and reuse in the cement industry: from the lab to the plant” in order to highlight the wide range of aspects covered by this event: European context regarding CCU, industrial CCU projects in the cement industry and also academic studies, especially those of the ECRA Chair carried out at UMONS. A “Save the date” email is under preparation and will be sent to potential participants during the next weeks.


8. REFERENCES

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9. ANNEXES

- Annex 1: solvents studied in the PhD Thesis of Sinda Laribi.

- Annex 2: draft Program of the Second ECRA Chair Scientific Event.


Annex 1

Conventional solvents like MEA 30 wt%, DEA 30 wt% …(see Table 1).

Table 1: Amine solvents

Chemical solvents

Product Name Manufacturer Purity Quality Developed formula

MEA Monoethanolamine Merck ≥ 99% For synthesis

MMEA Monomethylethanolamine Merck ≥ 98% For synthesis

DEA Diethanolamine Merck ≥ 99% For synthesis

MDEA N-methyldiethanolamine Merck ≥ 98% For synthesis

AMP 2-amino-2-methyl-1-

propanol Merck ~95% For synthesis

AHPD 2-amino-2–hydroxymethyl-

1,3-propanediol

VWR (BDH

PROLABO) 100%

AnalaR

NORMAPUR

Activated solvents using a blend of [amine+ activator] like [MEA 30% + TETRA 5%] or [AMP 30% + PZ 5%] using two types of activators PZ and TETRA. Likewise, the effect of PZ 10% and TETRA 30% is studied separately (see Table 2).

Table 2: Activator used in the solvents [amine+activator] Chemical solvents

Product Name Manufacturer Purity Quality Developed formula

PZ Piperazine Merck ≥ 99% For synthesis

TETRA Triethylenetetramine Merck ~95% For synthesis

All the scrubbing solvents are aqueous.


Annex 2

Draft Program of the Second ECRA Chair Scientific Event CO2 capture and reuse in the cement industry:

“From the lab to the plant” 09th November 2016 (Lixhe Cement plant visit on 10th November 2016)

09:00: Welcome coffee and registration

09:30: Introduction by Dean Pierre Dehombreux or Rector Calogero Conti? Faculty of Engineering – UMONS (B)

09:45: Presentation of the event and of the ECRA Chair activities by Prof Diane Thomas, Faculty of Engineering – UMONS (B)

10:00: European Context regarding Carbon Capture Storage & Utilization (CCSU) by European Commission representative?

10:20: Status of CCSU deployment in the industry against the background of roadmaps and industry’s challenges, by IEA (or IEAGHG) representative?

10:40: ECRA CCS Project: Status of the pilot plant preparation by Dr Martin Schneider, ECRA (D)

11:00: HeidelbergCement’s activities regarding Carbon Capture & Reuse by Daniel Gauthier, HeidelbergCement, Technical Advisory Board ECRA (B)

11:20– 11:50: Coffee break

11:50: LafargeHolcim’s activities: CCS and low CO2 cements (Solida), by M. Gimenez, LafargeHolcim

12:10: Cemcap Oxyfuel (preliminary prototype test results of oxyfuel calciner, burner and cooler),

by VDZ or any CemCap project representative 12:30: Integration of Ca-Looping systems (Cemcap),

by Polimi (I) (Advancements since 2014 by Prof. Consonni ?) or CSIC (S)?

13:00 – 14:15: Lunch and networking

14:15: Presentation of the Lixhe Cement plant and Leilac Project? by Julien Wart/Jan Theulen, HeidelbergCement (B)

14:35: ECRA Chair projects at UMONS: CO2 capture, purification and conversion into methanol and other materials by: Lionel Dubois (10 min) for global introduction on ECRA Chair Scientific Activities; Seloua Mouhoubi (20 min) on Post-Combustion CO2 capture by Demixing solvents, Sinda Laribi (20 min) on CO2 purification; Remi Chauvy (20 min) on CO2 reuse possibilities and Nicolas Meunier (20 min) on CO2 conversion into methanol, Post-doc and PhD Students ECRA Chair, Faculty of Engineering – UMONS (B) (+ 10 min questions)

16:15: CO2 valorization/conversion – Industrial case by ….

16:40: Closing of the day and acknowledgments: D. Thomas and M. Schneider Drink

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