Rémi Demol

Anthony Dufour

Yann Rogaume

Guillain Mauviel
email: guillain.mauviel@univ-lorraine.fr

Production of purified H 2 , heat and bio-char from wood: comparison between gasification and auto-thermal pyrolysis based on advanced process modeling

Biomass gasification is an interesting route for renewable hydrogen production, but it is still hampered by technical, environmental and economic issues. A first key step toward its development is the quantification of mass and energy balances of the integrated process. This work compares different processes to produce a purified H 2 from wood but also other products (heat, bio-char) at medium scale power (20 MW of biomass power inlet that corresponds to 3.7 t dry /h). Three complementary processes were modeled under Aspen Plus including biomass drying, gasification-pyrolysis reactors and advanced syngas upgrading units. The first two cases are based on oxygen/steam gasification 1) with or 2) without catalytic reactors (steam reforming and water gas-shift). The third case is an autothermal

oxidative pyrolysis resulting in bio-char and syngas. All the syngas cleaning process was detailed with a special focus on a partial oxidation (POX) unit to reduce the tar content. This unit was modeled by coupling Aspen Plus with Chemkin to predict tar and syngas composition by detailed elementary mechanisms. A hybrid hydrogen separation process is proposed combining membrane and pressure swing adsorption. A cape-open module for membrane modeling (called Memsic) was included in the whole process model. The global energetic efficiency is 75.4, 77.8 and 80.4%net for scenarios 1, 2 and 3, respectively. The hydrogen yields are 79, 26 and 18 g H 2 /kg biomass,dry after separation and heat efficiencies -corresponding to hot water production -were 23.4, 60.0 and 49.0%net respectively. The option 3 produces 110 g biochar /kg biomass,dry which is a carbon sink. All utilities and consumables were also determined. This model can be used for techno-economic and life cycle assessment studies. This methodology is also of interest to model all other thermochemical processes with detailed kinetics embedded in process models.

INTRODUCTION

Biomass is used for centuries to fulfill the heat and material demand for human activities. It is historically and even nowadays the first renewable energy [START_REF] Smil | Energy Transitions: Global and National Perspectives[END_REF][START_REF][END_REF] . The growing concerns about climate change resulting from anthropogenic emissions and the fore casted peak in oil production pushes researchers to develop innovative processes for the production of energy commodities from renewables. The taxation of greenhouse gas emissions is expected to favor the development of carbonneutral or even carbon-negative processes [START_REF] Dufour | Geological Sequestration of Biomass Char to Mitigate Climate Change[END_REF]4 .

Among the ways to decarbonize human activities, hydrogen knows a growing interest for mobility applications and for lowering industry carbon intensity 5 . Yet, even if this fuel does not release carbon dioxide during its combustion, its production-mainly from steam methane reforming, does. If H 2 is produced from water electrolysis, the CO 2 emission problem is then related to the electricity generation processes. Different policies tend to promote renewable hydrogen. France fixed the objective to increase the share of renewable hydrogen to 20-40% in 2028 in the industry sector [START_REF]Ministère de la Transition Energétique et Solidaire[END_REF] .

Besides water electrolysis from renewable electricity, the thermochemical processes also provide a potential way to produce hydrogen from biomass.

The gasification is the partial oxidation of a solid feedstock to produce a synthetic gas (syngas) made of CO, H 2 , CO 2 , CH 4 , H 2 O and light hydrocarbons. Polycyclic aromatic hydrocarbons (PAHs), NH 3 , HCl, H 2 S are also formed and must be removed from the syngas before its upgrading [START_REF] Woolcock | A Review of Cleaning Technologies for Biomass-Derived Syngas[END_REF] . The gasifying agent could be steam in order to maximize the concentration of hydrogen but it leads to highly endothermic behavior. The use of air leads to syngas diluted with nitrogen when the gasification is direct: this option should be avoided because it hampers the H 2 purification. To overcome this problem, a dual fluidized bed can be used [START_REF] Xu | The Superior Technical Choice for Dual Fluidized Bed Gasification[END_REF][START_REF] Corella | A Review on Dual Fluidized-Bed Biomass Gasifiers[END_REF] but it results in a complex technology notably due to loop seals [START_REF] Larsson | Experimental and Numerical Investigation of the Dynamics of Loop Seals in a Large-Scale DFB System under Hot Conditions[END_REF] . Another way is to use pure oxygen and steam instead of air [START_REF] Corella | 140gH2/Kg Biomass d.a.f. by a CO-Shift Reactor Downstream from a FB Biomass Gasifier and a Catalytic Steam Reformer[END_REF] . In any cases, the fluidized bed temperature is typically higher than 750°C to convert the pyrolysis char into syngas [START_REF] Bates | Modeling of Biomass Char Gasification, Combustion, and Attrition Kinetics in Fluidized Beds[END_REF] .

Another thermochemical process can produce hydrogen: the pyrolysis, that may also be achieved in fluidized bed [START_REF] Kersten | Biomass Pyrolysis in a Fluidized Bed Reactor. Part 1: Literature Review and Model Simulations[END_REF][START_REF] Bruchmüller | Modeling the Thermochemical Degradation of Biomass inside a Fast Pyrolysis Fluidized Bed Reactor[END_REF] . This process produces char, tar and permanent gas. To overcome the endothermicity of pyrolysis, a small amount of oxygen can be added to reach autothermal conditions [START_REF] Brown | Heterodoxy in Fast Pyrolysis of Biomass[END_REF] . But this pyrolysis step alone does not lead to high H 2 yields. In this article, it is proposed to do the partial oxidation of the tar and gas produced by pyrolysis in a downstream gas-phase reactor. This second step is achieved by mixing oxygen with pyrolysis gas in a partial oxidation (POX) reactor to reach high temperature (> 1000°C), thus producing a H 2 -rich syngas. Furthermore, the bio-char produced in the pyrolysis reactor enables carbon sequestration [START_REF] Dufour | Geological Sequestration of Biomass Char to Mitigate Climate Change[END_REF][START_REF] Lehmann | A Handful of Carbon[END_REF] .

The hydrogen in the syngas is relatively diluted (even with oxy-steam gasification) in the range 30-45%vol on a dry basis [START_REF] Corella | 140gH2/Kg Biomass d.a.f. by a CO-Shift Reactor Downstream from a FB Biomass Gasifier and a Catalytic Steam Reformer[END_REF][START_REF] Schmid | Product Gas Composition for Steam-Oxygen Fluidized Bed Gasification of Dried Sewage Sludge, Straw Pellets and Wood Pellets and the Influence of Limestone as Bed Material[END_REF][START_REF] Barisano | Steam/Oxygen Biomass Gasification at Pilot Scale in an Internally Circulating Bubbling Fluidized Bed Reactor[END_REF] . The production of hydrogen at high purity is difficult at this concentration for standard separation unit. Pressure swing adsorption (PSA H 2 ) is the classical technology. More than 70%vol of H 2 are required at the inlet [START_REF] Swanson | Techno-Economic Analysis of Biofuels Production Based on Gasification[END_REF] . To increase the content of H 2 , a reformer and water-gas-shift catalytic reactors should be added [START_REF] Corella | 140gH2/Kg Biomass d.a.f. by a CO-Shift Reactor Downstream from a FB Biomass Gasifier and a Catalytic Steam Reformer[END_REF][START_REF] Spath | Biomass to Hydrogen Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory Indirectly-Heated Gasifier[END_REF][START_REF] Fail | Wood Gas Processing To Generate Pure Hydrogen Suitable for PEM Fuel Cells[END_REF][START_REF] Kraussler | -h Long Term Operation of a Water Gas Shift Pilot Plant Processing Tar-Rich Product Gas from an Industrial Scale Dual Fluidized Bed Biomass Steam Gasification Plant[END_REF] . These reactors can be positioned after cold syngas cleaning 4,[START_REF] Spath | Biomass to Hydrogen Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory Indirectly-Heated Gasifier[END_REF][START_REF] Hamelinck | Future Prospects for Production of Methanol and Hydrogen from Biomass[END_REF] , or rather, downstream the gasifier to promote heat integration. The use of catalytic reactors after a gasifier has been demonstrated with catalytic reformers [START_REF] Caballero | Commercial Steam Reforming Catalysts To Improve Biomass Gasification with Steam-Oxygen Mixtures. 1. Hot Gas Upgrading by the Catalytic Reactor[END_REF][START_REF] Aznar | Commercial Steam Reforming Catalysts To Improve Biomass Gasification with Steam-Oxygen Mixtures. 2[END_REF][START_REF] Corella | 140gH2/Kg Biomass d.a.f. by a CO-Shift Reactor Downstream from a FB Biomass Gasifier and a Catalytic Steam Reformer[END_REF] and CO-shift catalytic reactors [START_REF] Corella | 140gH2/Kg Biomass d.a.f. by a CO-Shift Reactor Downstream from a FB Biomass Gasifier and a Catalytic Steam Reformer[END_REF][START_REF] Fail | Wood Gas Processing To Generate Pure Hydrogen Suitable for PEM Fuel Cells[END_REF][START_REF] Kraussler | -h Long Term Operation of a Water Gas Shift Pilot Plant Processing Tar-Rich Product Gas from an Industrial Scale Dual Fluidized Bed Biomass Steam Gasification Plant[END_REF][START_REF] Chianese | Experimental Investigations of Hydrogen Production from CO Catalytic Conversion of Tar Rich Syngas by Biomass Gasification[END_REF] . The tar content must be reduced below 2 g/Nm 3 dry basis (including benzene) at high temperature to avoid catalyst deactivation [START_REF] Aznar | Commercial Steam Reforming Catalysts To Improve Biomass Gasification with Steam-Oxygen Mixtures. 2[END_REF] . No deactivation from H 2 S was observed below 100 ppm [START_REF] Corella | 140gH2/Kg Biomass d.a.f. by a CO-Shift Reactor Downstream from a FB Biomass Gasifier and a Catalytic Steam Reformer[END_REF] .

The harvesting area of the biomass should be limited to minimize the economic and environmental impacts of its transport. Besides another argument for relatively small-scale biomass conversion process is linked to the fact that heat produced by the process should be valorized locally in order to increase the global efficiency [START_REF] François | Modeling of a Biomass Gasification CHP Plant: Influence of Various Parameters on Energetic and Exergetic Efficiencies[END_REF] . It is clearly easier to find a location for biomass gasifiers with a small heat demand (few MW) than a large one (dozens of MW). Finally, a local production of H 2 might be preferred for more direct and decentralized H 2 station for transport or industrial sectors, instead of a centralized production with H 2 transport by trucks [START_REF] Emonts | Flexible Sector Coupling with Hydrogen: A Climate-Friendly Fuel Supply for Road Transport[END_REF] .

The whole pyrogasification process must be modeled, from biomass drying to hydrogen separation, in order to assess the potential of the production of hydrogen from biomass. The study must also include co-products recovery and waste treatment. The modeling of pyrogasification processes has already been conducted, especially for the cleaning and conditioning of the syngas [START_REF] François | Modeling of a Biomass Gasification CHP Plant: Influence of Various Parameters on Energetic and Exergetic Efficiencies[END_REF][START_REF] Francois | Detailed Process Modeling of a Wood Gasification Combined Heat and Power Plant[END_REF] and also for hydrogen production at large-scale facilities [START_REF] Swanson | Techno-Economic Analysis of Biofuels Production Based on Gasification[END_REF][START_REF] Spath | Biomass to Hydrogen Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory Indirectly-Heated Gasifier[END_REF] . Gasification is often modeled as a combination of RYIELD and RGIBBS reactors. Firstly, the biomass is decomposed into its elemental stable components (H 2 , C, O 2 , N 2 , H 2 S, HCl). Then the RGIBBS reactor estimates the equilibrium composition at a given temperature [START_REF] Puig-Arnavat | Review and Analysis of Biomass Gasification Models[END_REF] .

This model can give a rough estimate of the main components but it is unable to predict the yields of minor products (tars) which are the bottleneck of gasification. The tar formation and up-grading has been modeled but the chosen models are frequently overly simplistic [START_REF] Srinivas | Modeling Tar Handling Options in Biomass Gasification[END_REF] . Some research groups developed a fluidized bed model to predict the main products and some secondary products [START_REF] Abdelouahed | Detailed Modeling of Biomass Gasification in Dual Fluidized Bed Reactors under Aspen Plus[END_REF][START_REF] Gómez-Barea | Modeling of Biomass Gasification in Fluidized Bed[END_REF][START_REF] Bates | Steam-Air Blown Bubbling Fluidized Bed Biomass Gasification (BFBBG): Multi-Scale Models and Experimental Validation[END_REF] .

To the best of our knowledge, advanced models of the complete process, from biomass to purified H 2 , are still scarce, notably if one considers tar formation and upgrading.

Spath et al. has studied in a pioneering work the modeling of the complete process of hydrogen production from biomass gasification in an indirectly-heated gasifier. The steam reformer and water gas-shift reactors was positioned after wet scrubbing of tars inducing a heat penalty on the process [START_REF] Spath | Biomass to Hydrogen Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory Indirectly-Heated Gasifier[END_REF] .

The scale (2000 dry ton/day) of this process makes possible the use of catalytic reactors but it requires long-distance collection of biomass. They used empirical correlations to model syngas and tar composition.

Martín and Grossmann presented the basis of a superstructure optimization for the production of Fischer-Tropsch diesel from biomass [START_REF] Martín | Process Optimization of FT-Diesel Production from Lignocellulosic Switchgrass[END_REF] . The optimal solution to reach the targeted CO/H 2 ratio was composed of an indirect gasifier and a steam reformer instead of direct gasification coupled to partial oxidation. No further composition adjustment (waster gas-shift, PSA H 2 ) was necessary for this application [START_REF] Martín | Process Optimization of FT-Diesel Production from Lignocellulosic Switchgrass[END_REF] .

Syngas cleaning and upgrading processes usually consist in tar reformer, water gas-shit reactor and PSA H 2 [START_REF] Ersöz | Investigation of a Novel & Integrated Simulation Model for Hydrogen Production from Lignocellulosic Biomass[END_REF][START_REF] Gupta | Hydrogen from Biomass by Oxy-Steam Gasification -A Quantitative Analysis of Cases[END_REF][START_REF] Kalinci | Exergoeconomic Analysis of Hydrogen Production from Biomass Gasification[END_REF][START_REF] Ishaq | A Novel Biomass Gasification Based Cascaded Hydrogen and Ammonia Synthesis System Using Stoichiometric and Gibbs Reactors[END_REF][START_REF] Koroneos | Hydrogen Production via Biomass Gasification -A Life Cycle Assessment Approach[END_REF] , possibly with Sulphur removal with a chelated iron solution (LO-CAT process) [START_REF] Susmozas | Life-Cycle Performance of Hydrogen Production via Indirect Biomass Gasification with CO2 Capture[END_REF] . The tar reformer can be replaced by catalytic filter candles [START_REF] Pallozzi | Performance Evaluation at Different Process Parameters of an Innovative Prototype of Biomass Gasification System Aimed to Hydrogen Production[END_REF] . When tars species were considered, only few surrogate molecules were included [START_REF] Ersöz | Investigation of a Novel & Integrated Simulation Model for Hydrogen Production from Lignocellulosic Biomass[END_REF][START_REF] Pallozzi | Performance Evaluation at Different Process Parameters of an Innovative Prototype of Biomass Gasification System Aimed to Hydrogen Production[END_REF] . The purified H 2 yield was estimated to 76.1 [START_REF] Ersöz | Investigation of a Novel & Integrated Simulation Model for Hydrogen Production from Lignocellulosic Biomass[END_REF] , 55.0 41 , 75.2 [START_REF] Pallozzi | Performance Evaluation at Different Process Parameters of an Innovative Prototype of Biomass Gasification System Aimed to Hydrogen Production[END_REF] or 107.4 37 g H2 /kg biomass . On an energetic basis, Kalinci et al. showed that gasifier and PSA exhibit the most energy and exergy losses along the process [START_REF] Kalinci | Exergoeconomic Analysis of Hydrogen Production from Biomass Gasification[END_REF] .

Marcantonio et al. modeled a circulating bubbling fluidized bed with a quasi-equilibrium approach model validated on experimental data from a pilot plant. They also investigated the use of a palladium membrane that gave a better H 2 recovery [START_REF] Marcantonio | Process Analysis of Hydrogen Production from Biomass Gasificati on in Fluidized Bed Reactor with Different Separation Systems[END_REF] .

The purification of the syngas was investigated in more details with Ribeiro et al . by modeling the detailed PSA cycle to remove CO 2 from the syngas. H 2 and CO were dedicated for Fischer-Tropsch fuels production [START_REF] Ribeiro | Pressure Swing Adsorption for CO2 Capture in Fischer-Tropsch Fuels Production from Biomass[END_REF] . To the best of our knowledge, no model was published on membrane combined with PSA for H 2 separation.

Our research group has developed previous Aspen Plus® models about biomass gasification and oxidation [START_REF] François | Modeling of a Biomass Gasification CHP Plant: Influence of Various Parameters on Energetic and Exergetic Efficiencies[END_REF][START_REF] Francois | Detailed Process Modeling of a Wood Gasification Combined Heat and Power Plant[END_REF][START_REF] Abdelouahed | Detailed Modeling of Biomass Gasification in Dual Fluidized Bed Reactors under Aspen Plus[END_REF]45,[START_REF] François | Assessing the Fate of Nutrients and Carbon in the Bioenergy Chain through the Modeling of Biomass Growth and Conversion[END_REF] for heat or power production. Here, we complete our previous work on different pyro-gasification processes dedicated to the production of purified hydrogen and bio-char, with different gasification reactors and syngas refining units. We have also improved our modeling approach by including detailed kinetic mechanisms embedded under Aspen Plus.

The aim of this work is to provide detailed mass and energy balance s for three processes along with utilities and consumables. These data are essential for further techno-economic and environmental assessment.

The first scenario considers the maximization of hydrogen production and a residual heat production.

The second one is a simpler and probably cheaper process with lower hydrogen production but higher heat production. The third case is based on oxidative pyrolysis to produce bio-char (carbon sink), hydrogen and heat. All these options were modeled in Aspen Plus® associated with experimental data obtained from the literature and from a semi-industrial pilot plant (University of Lorraine, Epinal, France) [START_REF] Debal | Wood Gasification in a Semi-Industrial Bubbling Fluidized Bed Gasifier[END_REF] .

Therefore, the novelty of this work can be outlined by these three main aspects:

1) To the best of our knowledge, these three processes were not yet modeled with the proposed detailed approach under the Aspen Plus framework, including elementary reactions for gasphase reactions, hydrodynamic of fluidized bed, and advanced purification of H 2 (membrane permeation and PSA adsorption).

2) Novel results on gas cleaning in a Venturi scrubber on a gasification pilot plant are presented and embedded in the Aspen Plus model.

3) These three main routes of H 2 , heat and bio-char production are compared and discussed based on their energy, mass and hydrogen balance.

MODEL DESCRIPTION

Scenarios investigated

Three scenarios of hydrogen production from biomass were investigated (Figure 1). The three scenarios were designed for a territorial scale of 20 MW LHV-basis (30 kt dry biomass/year). The gasifier is a single fluidized bed reactor, which is cheaper and simpler to manage at territorial scale in comparison to dual fluidized bed technology. The gasifying agent is a mixture of oxygen and steam.

Case 1 aims at maximizing the production of H 2 by implementing a steam reformer and water-gas shift catalytic reactors downstream the fluidized bed. It is based on experimental results of Corella et al. [START_REF] Corella | 140gH2/Kg Biomass d.a.f. by a CO-Shift Reactor Downstream from a FB Biomass Gasifier and a Catalytic Steam Reformer[END_REF] .

Case 2 considers a simpler process without catalytic reactors, targeting a lower production of H 2 and a higher production of heat. The last case investigates the autothermal pyrolysis of biomass to produce bio-char and a H 2 -rich gas obtained after the partial oxidation of the pyrolysis gas.

Description of the processes and modeling

The modeling approach is presented in Figure 1. The main assumptions of each process units are summarized in Table 2 and discussed here after.

The feedstock considered was woodchips, a by-product of the forest harvesting and wood industry. This biomass was supposed to be crushed directly in the forest or in dedicated platforms. The humidity was fixed to 40%wt after delivery. Table S1 provides the detailed composition of the feedstock.

The detailed flow sheets of the three options are presented in supporting information (SI) (Figure S1-S3). The equation of state RK-Aspen was used as it is recommended for hydrocarbon mixtures and light gases 48 . The species BIOMASS and CHAR were modeled as non-conventional solids with their proximate and elemental composition. The heat of combustion was estimated with Mott and Spooner model, which is tailored for biomass and its high oxygen content [START_REF] Rönsch | Calculation of Heating Values for the Simulation of Thermo-Chemical Conversion Plants with Aspen Plus[END_REF] .

Biomass drying

The biomass was considered dried down to 20%wt with the low-temperature heat contained in the boiler exhaust gas in order to increase energy efficiency and to limit the amount of tars produced by the gasifier [START_REF] Corella | Calculation of the Conditions to Get Less than 2 g Tar/Mn3 in a Fluidized Bed Biomass Gasifier[END_REF] . The dryer model was taken from François et al. [START_REF] François | Modeling of a Biomass Gasification CHP Plant: Influence of Various Parameters on Energetic and Exergetic Efficiencies[END_REF] . It estimates the VOCs emissions during the drying.

Oxygen production

Concerning the production of oxygen as gasifying agent, the VSA O 2 (vacuum swing adsorption) process is the most adapted one for small-scale production in the range of 10 to 200 tons of O 2 per day and if very highly pure O 2 is not required (93-95%v, the rest is mainly argon) [START_REF] Smith | A Review of Air Separation Technologies and Their Integration with Energy Conversion Processes[END_REF] . The VSA O 2 was modeled as a simple separator to reach a purity of 93%v and a recovery rate of 55% [START_REF] Ding | Optimization and Analysis of the VPSA Process for Industrial-Scale Oxygen Production[END_REF] . Air was compressed at 1.5 bar before the columns. The purge pressure was set to 0.6 bar obtained with a vacuum pump. An adsorbent commonly used consists of lithium-doped zeolites. The required adsorbent quantity was estimated from ref. [START_REF] Silva Ferreira | High-Purity Oxygen Production by VPSA[END_REF] .

Gasifier

The gasification and the pyrolysis were conducted in a bubbling fluidized bed, which is the most adapted technology for the targeted scale [START_REF] Dufour | Optimisation de la production d'hydrogène par conversion du méthane dans les procédés de pyrolyse/gazéification de la biomasse[END_REF] .

For cases 1 and 2, a real syngas composition from literature data was used in order to have a detailed and accurate composition of tars. Among few detailed results available in literature [START_REF] Schmid | Product Gas Composition for Steam-Oxygen Fluidized Bed Gasification of Dried Sewage Sludge, Straw Pellets and Wood Pellets and the Influence of Limestone as Bed Material[END_REF][START_REF] Barisano | Steam/Oxygen Biomass Gasification at Pilot Scale in an Internally Circulating Bubbling Fluidized Bed Reactor[END_REF] , the experimental results obtained by Schmid et al. [START_REF] Schmid | Product Gas Composition for Steam-Oxygen Fluidized Bed Gasification of Dried Sewage Sludge, Straw Pellets and Wood Pellets and the Influence of Limestone as Bed Material[END_REF] in a steam/oxygen fluidized bed were selected (see SI S2). Their operating conditions were tailored to the production of hydrogen with an equivalent ratio (ER) of 0.25 and a molar steam to carbon ratio of 1 for a bed temperature at 850°C. The solid organic residue (char) yield was estimated to 10 g daf/kg of dry biomass. Gil et al. mentioned 5-20 g/kg daf as char yield for a steam-oxygen bubbling fluidized bed [START_REF] Gil | Biomass Gasification in Atmospheric and Bubbling Fluidized Bed: Effect of the Type of Gasifying Agent on the Product Distribution[END_REF] . Therefore, 10 g daf/kg dry biomass of char is an average common value for char yield produced by this technology. The elemental composition of char was assumed as: 85% C, 2% H and 13% O daf. The global solid residue recovered is made up of char and ash. An external Fortran subroutine linked to RYIELD model was used to compute the gas and tar composition according to the experimental results and to the biomass flow rate. Atomic balances in C, H, O, Cl and S were ensured by adjusting the CO 2 , H 2 O, O 2 , HCl and H 2 S flows, respectively. The heat balance was used in the RYIELD model to calculate the temperature of the syngas at the outlet assuming an adiabatic reactor.

Pyrolyser

In the case 3, an adiabatic fluidized bed pyrolyser was used. A small amount of oxygen was injected (auto-thermal conditions) to provide heat internally [START_REF] Brown | Heterodoxy in Fast Pyrolysis of Biomass[END_REF] . As a consequence, its behavior was very close to the auto-thermal fluidized bed gasifier used in the first scenarios. The main difference was the bed temperature below 600°C instead of 850°C. In this condition, pyrolysis char was an important product to be recovered. To the best of the author's knowledge, no detailed data on the auto-thermal pyrolysis presenting gas molecular composition (gas and tar) is available in the literature. Therefore, in order to model this auto-thermal pyrolysis, a model was developed to estimate the yield and detailed composition of pyrolysis products (char, gas, water and tars). Ranzi's model of biomass pyrolysis [START_REF] Debiagi | Detailed Kinetic Mechanism of Gas-Phase Reactions of Volatiles Released from Biomass Pyrolysis[END_REF] and radical kinetic mechanisms [START_REF] Dhahak | Development of a Detailed Kinetic Model for the Combustion of Biomass[END_REF][START_REF] Norinaga | Analysis of Pyrolysis Products from Light Hydrocarbons and Kinetic Modeling for Growth of Polycyclic Aromatic Hydrocarbons with Detailed Chemistry[END_REF][START_REF] Norinaga | Detailed Chemical Kinetic Modelling of Vapour-Phase Cracking of Multi-Component Molecular Mixtures Derived from the Fast Pyrolysis of Cellulose[END_REF] were used. The ER used was set to 0.10, slightly higher than the minimum ER of 0.08 estimated by Brown for autothermal pyrolysis at around 500°C to compensate heat losses [START_REF] Brown | Heterodoxy in Fast Pyrolysis of Biomass[END_REF] . The heat balance showed that an ER equals to 0.10 was able to reach a mean temperature in the fluidized bed of 565°C. Additional information on the pyrolysis model and a comparison with experimental results can be found in SI S3. A part of the syngas was recycled to the fluidized bed to maintain a fluidization velocity consistent to the gasification cases. An external Fortran subroutine was used to ensure mass balance similarly to the gasification reactor.

Partial oxidation, steam reformer and water-gas shift

When the production of H 2 is maximized (case 1), the process includes catalytic steam reformer and water gas shift units. These two steps were conducted at high temperature after the gasifier in order to promote heat integration and according to the experiments of Corella et al. [START_REF] Corella | 140gH2/Kg Biomass d.a.f. by a CO-Shift Reactor Downstream from a FB Biomass Gasifier and a Catalytic Steam Reformer[END_REF] . Unfortunately, the catalysts are sensitive to the concentration of tars which may cause their deactivation. 2 g/Nm [START_REF] Dufour | Geological Sequestration of Biomass Char to Mitigate Climate Change[END_REF] (including benzene, dry basis) was recommended by Corella et al. as the targeted tar content for maintaining the stability of the catalytic reformer [START_REF] Aznar | Commercial Steam Reforming Catalysts To Improve Biomass Gasification with Steam-Oxygen Mixtures. 2[END_REF] . Therefore, a partial oxidizing unit was used after the gasifier to reduce the tar content down to 2 g/Nm [START_REF] Dufour | Geological Sequestration of Biomass Char to Mitigate Climate Change[END_REF] . This limit can also be obtained by an optimized design and operation of the gasifier even if it was not the case with the experimental data used here for syngas composition since olivine was used as bed material. A better catalyst (dolomite, nickelolivine) can contribute to the reduction of tar content, but it would also increase the operating costs.

The addition of a small amount of oxygen in the syngas leads to the oxidation and cracking of tars at high temperature (over 1000°C). This POX unit was modeled by detailed kinetic models 57 -59 which were implemented by coupling Aspen Plus with ANSYS Chemkin Pro 17.0 (SI S4). For the gasification scenarios (1-2), the oxygen was adjusted to an equivalent ratio of 0.12 to reach the target of 2 g/Nm [START_REF] Dufour | Geological Sequestration of Biomass Char to Mitigate Climate Change[END_REF] of tars. More details on the impact of this ER are given in SI, section S5.

Experiments on steam reforming and water-gas-shift with a real syngas were conducted in the literature [START_REF] Aznar | Commercial Steam Reforming Catalysts To Improve Biomass Gasification with Steam-Oxygen Mixtures. 2[END_REF][START_REF] Kraussler | Performance of a Water Gas Shift Pilot Plant Processing Product Gas from an Industrial Scale Biomass Steam Gasification Plant[END_REF] . The reformer reactor was modeled with RPLUG and the kinetics of Corella's team [START_REF] Aznar | Commercial Steam Reforming Catalysts To Improve Biomass Gasification with Steam-Oxygen Mixtures. 2[END_REF] with a nickel-based catalyst (reactions and kinetics presented in SI S6). The dimensions of the reactor were adjusted to reach 95% conversion of methane according to Caballero et al. [START_REF] Caballero | Commercial Steam Reforming Catalysts To Improve Biomass Gasification with Steam-Oxygen Mixtures. 1. Hot Gas Upgrading by the Catalytic Reactor[END_REF] . This kinetic approach allowed to predict the remaining tars after the catalytic reformer.

Then, two stages of water gas shift were used, first at high temperature (350°C), second at a lower temperature (200°C) to promote CO conversion. The catalysts commonly used are iron/chromium oxide and Cu/ZnO/Al 2 O 3 for high and low temperature respectively [START_REF] Corella | 140gH2/Kg Biomass d.a.f. by a CO-Shift Reactor Downstream from a FB Biomass Gasifier and a Catalytic Steam Reformer[END_REF] . The water-gas-shift reactors were modeled with RGIBBS model and a temperature approach of 20°C [START_REF] Spath | Biomass to Hydrogen Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory Indirectly-Heated Gasifier[END_REF] . A steam to CO ratio of 3 was used to maximize the H 2 production.

Wet scrubbing of syngas

As a final syngas polishing, water scrubber was chosen to remove residual tars and other contaminants (NH 3 , HCl, H 2 S). Even if this operation was not required for tar removal in case 1, the wet scrubber has another purpose: the condensation of the syngas water content. Table 1 presents the tar removal efficiency found in literature and based on a pilot system experiment with Venturi and wet scrubbers in series. This pilot plant at University of Lorraine (Epinal, France) can operate 50 kg biomass/h. The scrubbing water flow rate is about 1 m 3 /h. This step was modeled as a FLASH unit and the composition of tars adjusted accordingly to the experimental results (of Table 1). The removal of NH 3 , H 2 S and HCl was modeled with ELECNRTL model [START_REF] François | Modeling of a Biomass Gasification CHP Plant: Influence of Various Parameters on Energetic and Exergetic Efficiencies[END_REF] .

Wastewater treatment

The species removed from the syngas and present in the scrubbing water were separated by coagulation and flotation in a dissolved air flotation unit (DAF) with the addition of soda to increase the pH along with flocculants and coagulant. The excess of water resulting from condensing water from the syngas was removed and sent to district water system after a fixed bed of activated carbon to remove the residual contaminants [START_REF] Yousef | Adsorption as a Process for Produced Water Treatment: A Review[END_REF][START_REF] Yuan | Adsorption of Polycyclic Aromatic Hydrocarbons from Water Using Petroleum Coke-Derived Porous Carbon[END_REF] . The amount of activated carbon was estimated with ref. [START_REF] Yuan | Adsorption of Polycyclic Aromatic Hydrocarbons from Water Using Petroleum Coke-Derived Porous Carbon[END_REF] . Syngas temperature around 150°C and 30°C at the inlet and outlet respectively, scrubbing water between 25 and 35°C at the inlet and outlet respectively. b ECN classification [START_REF]Tar Dew Point -Complete Model[END_REF] . c This value is an average that takes into account the relative yields of tars in this class.

Hydrogen separation

The standard process for H 2 separation is the pressure swing adsorption unit (PSA). Yet, the inlet concentration of H 2 should be at least 70%v to reach a high purity separation (99.99%v) [START_REF] Spath | Biomass to Hydrogen Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory Indirectly-Heated Gasifier[END_REF] . To achieve such high concentration at the inlet, a part of the pure hydrogen produced can be recycled at the inlet of the PSA [START_REF] Spath | Biomass to Hydrogen Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory Indirectly-Heated Gasifier[END_REF] . However, when the concentration of H 2 was too low (cases 2 and 3), a polyimide membrane permeable to H 2 was used before the PSA. The membrane plays the role of a H 2 preconcentrator (more details are given in SI S7).

The PSA was modeled as a SEP block with fixed recovery and purity, a part of the product was recycled to reach 70%v content in H 2 [START_REF] Spath | Biomass to Hydrogen Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory Indirectly-Heated Gasifier[END_REF] . The recovery rate was assumed to be 85%. The amount of adsorbent (zeolite and activated carbon) was estimated based on NREL calculation [START_REF] Golmakani | Investigating PSA, VSA, and TSA Methods in SMR Unit of Refineries for Hydrogen Production with Fuel Cell Specification[END_REF] . The membrane permeable to H 2 was modeled with the cape-open model called "MEMSIC" developed in our laboratory [START_REF] Bounaceur | Rigorous Variable Permeability Modelling and Process Simulation for the Design of Polymeric Membrane Gas Separation Units: MEMSIC Simulation Tool[END_REF] .

The optimal design of the hybrid H 2 separation unit was determined based on the specific separation cost. The method used for the determination of the optimal architecture (membrane surface, pressure on the retentate and permeate of the membrane) is described in SI S7. The goal was to achieve 99.99%vol hydrogen purity of hydrogen provided at 70 bar.

Heat generation

The tail gas of the hydrogen separation unit still contains some H 2 , CO, CH 4 and C 2 . Its lower heating value was too low for using it in an internal combustion engine for electricity production. Therefore, this gas was burnt in a gas boiler to produce heat for a heating network. The temperature range of the water network was 40/80°C. The gas boiler model was adapted from François et al. [START_REF] François | Modeling of a Biomass Gasification CHP Plant: Influence of Various Parameters on Energetic and Exergetic Efficiencies[END_REF] . We have implemented a Fortran subroutine fixing the pollutants yields based on the exhaust gas concentration of an industrial gas boiler. The atomic mass balance was computed with the same procedure as for the gasifier model.

Thermal integration

A pinch analysis was performed to build the heat exchanger network. The steam required for gasification was obtained with heat exchangers cooling the syngas before the reformer and the watergas shift (case 1) or before the wet scrubber (case 2). The excess of heat was recovered for the heating network. Syngas and hydrogen compression requires multistage compression with intercooler. A part of this heat was used for preheating steam flow to feed the gasifier and the steam reformer, another part was recovered for the heat network. A small amount of cold water (15°C) was required as cooling utility to reach the lowest temperatures level in the process (30°C, between two compression stages). Assuming SV 14 000 h -1 and density 1025 kg/m 3 24 , catalyst replacement 33% per year. f Assuming SV 2 700 h -1 for HTS (iron and chromium oxide BASF K6-11 in [START_REF] Larsson | Experimental and Numerical Investigation of the Dynamics of Loop Seals in a Large-Scale DFB System under Hot Conditions[END_REF] ) and SV 5100 h -1 for LTS (Cu/ZnO/Al 2 O 3 BASF K3-110 in [START_REF] Corella | 140gH2/Kg Biomass d.a.f. by a CO-Shift Reactor Downstream from a FB Biomass Gasifier and a Catalytic Steam Reformer[END_REF] ) and catalyst density 897.0 kg/m 3 , catalyst replacement 33% per year.

g Chemicals required for Dissolve Air Flotation unit and activated carbon guard bed for residual tars. Amount of activated carbon estimated assuming 3.25 g PAH adsorbed per g of activated carbon [START_REF] Yuan | Adsorption of Polycyclic Aromatic Hydrocarbons from Water Using Petroleum Coke-Derived Porous Carbon[END_REF] . h Mass of fresh adsorbent per year estimated from Peters et al. and Swanson et al. [START_REF] Swanson | Techno-Economic Analysis of Biofuels Production Based on Gasification[END_REF][START_REF] Peters | Plant Design and Economics for Chemical Engineers[END_REF] , adsorption isotherms for activated carbon and zeolite adsorbent [START_REF] Golmakani | Investigating PSA, VSA, and TSA Methods in SMR Unit of Refineries for Hydrogen Production with Fuel Cell Specification[END_REF] assuming a 4-year lifetime (7500 hours of operations). i Since membrane module lifetime is expected to last 5 years, it assumed 20% of membrane surface replacement per year.

Definition of energetic efficiency

The energetic efficiency is defined with reference to the lower heating value of woodchips on dry basis.

The net 𝜂 𝑛𝑒𝑡 and gross 𝜂 𝑔𝑟𝑜𝑠𝑠 energetic efficiencies are linked to heat 𝜂 ℎ𝑛 , hydrogen 𝜂 𝐻 2 and bio-char 𝜂 𝑏𝑖𝑜𝑐ℎ𝑎𝑟 efficiency.

𝜂 𝑛𝑒𝑡 = 𝜂 ℎ𝑛 + 𝜂 𝐻 2 + 𝜂 𝑏𝑖𝑜𝑐ℎ𝑎𝑟 -𝜂 𝑝𝑜𝑤𝑒𝑟 𝑐𝑜𝑛𝑠𝑢𝑚𝑝𝑡𝑖𝑜𝑛 (1)
𝜂 𝑔𝑟𝑜𝑠𝑠 = 𝜂 ℎ𝑛 + 𝜂 𝐻 2 + 𝜂 𝑏𝑖𝑜𝑐ℎ𝑎𝑟 (2)
𝜂 ℎ𝑛 = 𝑄 ̇ℎ𝑛 𝑚̇𝑤 𝑜𝑜𝑑,𝑑𝑟𝑦 • 𝐿𝐻𝑉 𝑤𝑜𝑜𝑑,𝑑𝑟𝑦 (3)
𝜂 𝐻 2 = 𝑚̇𝐻 2 • 𝐿𝐻𝑉 𝐻 2 𝑚̇𝑤 𝑜𝑜𝑑,𝑑𝑟𝑦 • 𝐿𝐻𝑉 𝑤𝑜𝑜𝑑,𝑑𝑟𝑦 (4)
𝜂 𝑏𝑖𝑜𝑐ℎ𝑎𝑟 = 𝑚̇𝑏 𝑖𝑜𝑐ℎ𝑎𝑟 • 𝐿𝐻𝑉 𝑏𝑖𝑜𝑐ℎ𝑎𝑟,𝑑𝑟𝑦 𝑚̇𝑤 𝑜𝑜𝑑,𝑑𝑟𝑦 • 𝐿𝐻𝑉 𝑤𝑜𝑜𝑑,𝑑𝑟𝑦 (5)
𝜂 𝑝𝑜𝑤𝑒𝑟 𝑐𝑜𝑛𝑠𝑢𝑚𝑝𝑡𝑖𝑜𝑛 = 𝑊 ̇𝑐𝑜𝑛𝑠𝑢𝑚𝑝𝑡𝑖𝑜𝑛 𝑚̇𝑤 𝑜𝑜𝑑,𝑑𝑟𝑦 • 𝐿𝐻𝑉 𝑤𝑜𝑜𝑑,𝑑𝑟𝑦 (6)
𝑄 ̇ℎ𝑛 is the heat power sent to the heat network, 𝑚̇𝑤 𝑜𝑜𝑑,𝑑𝑟𝑦 , 𝑚̇𝐻 2 and 𝑚̇𝑏 𝑖𝑜𝑐ℎ𝑎𝑟 the mass flowrate of biomass, hydrogen and biochar. 𝑊 ̇𝑐𝑜𝑛𝑠𝑢𝑚𝑝𝑡𝑖𝑜𝑛 is the electrical power consumption. 𝐿𝐻𝑉 𝑏𝑖𝑜𝑐ℎ𝑎𝑟,𝑑𝑟𝑦 and 𝐿𝐻𝑉 𝑤𝑜𝑜𝑑,𝑑𝑟𝑦 are the lower heating value of biochar and wood.

RESULTS AND DISCUSSION

Energy & Mass balance

The energy balances of each option are presented in Figure 3. The mass balance and the detailed composition of the main streams are available in SI S1. The main utilities consumptions including catalysts and adsorbents (Activated carbon AC and zeolite Ze) are given in Table 3. The electricity consumption was mainly driven by the hydrogen separation step and its compressors. The first option was the most electricity consuming because the entire syngas was compressed to 25 bar whereas the first stage of separation in cases 2 and 3 required a lower pressure (5 bar) (see supporting information S7). The electricity demand was also higher because of the final compression of H 2 on a bigger flow rate. The estimated amount of activated carbon to clean the excess water of residual PAHs and other contaminants was lower in the case 1 since the steam reformer reactor exhibits a catalytic effect on the tar reduction. The worst case for activated carbon consumption was the third case due to a higher quantity of tars in the syngas before wet scrubbing.

Spath et al. studied a similar process with indirectly heated biomass gasifier on a larger scale (434 MW LHV-basis). They determined a gross efficiency of 49.8% and a net efficiency of 45.6% for the production of hydrogen [START_REF] Spath | Biomass to Hydrogen Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory Indirectly-Heated Gasifier[END_REF] . We found in this work a higher H 2 efficiency (57.6% gross and 52.0% net in case 1). This is mainly due to a better conversion of biomass into H 2 and CO in our case. At the exit of the gasifier, the yields were 24.8 mol H 2 /kg biomass dry and 11.7 mol CO /kg biomass dry (O 2 /H 2 O bubbling fluidized bed [START_REF] Schmid | Product Gas Composition for Steam-Oxygen Fluidized Bed Gasification of Dried Sewage Sludge, Straw Pellets and Wood Pellets and the Influence of Limestone as Bed Material[END_REF] ) compared to 8.4 mol H 2 /kg biomass dry and 14.8 mol CO /kg biomass dry in Spath et al. [START_REF] Spath | Biomass to Hydrogen Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory Indirectly-Heated Gasifier[END_REF] (a dual fluidized bed).

The second case without catalytic reactors gives a similar global efficiency (77.8% net) but the production of heat was higher (60.0% instead of 23.4%) and the production of H 2 smaller (17.9% instead of 52.0%). The second case is less interesting based on H 2 production, but its CAPEX and OPEX are probably significantly lower (two catalytic reactors were removed). Furthermore, this process can present an interest for the co-production of H 2 and heat for sites with higher heat demands.

The third case of autothermal pyrolysis and its three products (hydrogen, heat and bio-char) gives a higher efficiency (80.4% net) when the bio-char is considered as an energy product. The two energy vectors (hydrogen and heat) represent a 51.2%net efficiency. A large amount of syngas was recycled to the pyrolyser to maintain its fluidization.

In the autothermal pyrolysis option (case 3), 408 kg/h of bio-char are produced corresponding to a char yield around 11.0%daf and a carbon yield of 16.6%. This bio-char yield should be considered with caution because it is estimated from Ranzi's model of biomass pyrolysis and not validated in a pilot

plant. Yet, this model gives a rough estimate of the bio-char composition and the gas and tar detailed composition. As a comparison, Polin et al. conducted autothermal pyrolysis at 500°C in a fluidized bed of two biomasses: Red Oak and corn stover with an ER of 0.10 and 0.068, respectively. The biochar yields were 9.5%wt for Red Oak and 20.1%wt for corn stover. The corresponding carbon yields were estimated to 14.5% for Red Oak and 26.7% for corn stover [START_REF] Polin | Process Intensification of Biomass Fast Pyrolysis through Autothermal Operation of a Fluidized Bed Re actor[END_REF][START_REF] Polin | Conventional and Autothermal Pyrolysis of Corn Stover: Overcoming the Processing Challenge s of High-Ash Agricultural Residues[END_REF] . The Red Oak experiments are compared with the results of this model in SI 3. The model predicts nicely the overall permanent gas mass yield but over-predicts the char yield (see SI 3 for more details).

Fate of hydrogen along the process unit

Figure 4 shows the molar flow rate and fraction of hydrogen along the process for each option. The objective was to reach 70%vol of H 2 before the PSA to produce quasi-pure hydrogen with the PSA H 2 . The flow rate of hydrogen is doubled when catalytic reformer and water-gas shift reactors are used (Figure 4). Its molar fraction is increased from 35.5% to 55.5%db. This concentration is reached in case 2 and 3 with a membrane module (67.3 and 66.8%db respectively for cases 2 and 3). The 70%vol concentration is achieved by recirculating a part of the hydrogen produced.

In the best case (1), 107 g H 2 /kg biomass,dry could be produced in which 79 g are effectively separated (76% recovery). This result compares well with the 140 g H 2 /kg biomass,dry produced claimed by Corella et al. [START_REF] Corella | 140gH2/Kg Biomass d.a.f. by a CO-Shift Reactor Downstream from a FB Biomass Gasifier and a Catalytic Steam Reformer[END_REF] for oxy-steam gasification followed by reformer and shift reactors. This lower yield of hydrogen can be explained by the partial oxidation of a part of the hydrogen in the POX unit. This reactor is necessary because the syngas produced by Schmid et al. [START_REF] Schmid | Product Gas Composition for Steam-Oxygen Fluidized Bed Gasification of Dried Sewage Sludge, Straw Pellets and Wood Pellets and the Influence of Limestone as Bed Material[END_REF] When catalytic reactors are removed, the hydrogen production is divided by a factor of 2 or 3. In case 2, 47 g H 2 /kg biomass,dry could be produced in which 26 g are effectively separated. In the case 3, 31 g H 2 /kg biomass,dry could be produced and 18 g are separated. These two cases present smaller hydrogen production but they could be more suitable for a territorial level: the processes are simpler and more robust than with catalytic reactors. The separation process of hydrogen was chosen to minimize the hydrogen specific separation cost but this architecture did not lead necessarily to the maximum hydrogen recovery rate. A higher production of H 2 would lead to higher specific separation cost. In case 3, char is produced and may be used to create a carbon sink.

The final H 2 yield of case study cannot be compared because our separation process had never been previously proposed.

In order to better understand the H atoms transfer from wood and water to H 2 , Figure 5 shows the fate of H along the process units. In the first scenario, 42.7% of the produced H 2 comes from the oxysteam gasification of biomass after the POX unit, 13.2% results from steam reforming and 44.2% from water gas-shift reactors. As a consequence of catalytic reactors, almost all the hydrogen content in the syngas before the separation process is attributed to the H 2 molecule. In the second case, all the H 2 comes from biomass and steam during gasification. As in case 1, after the partial oxidation of the syngas to reduce the amount of tars, the hydrogen yield was slightly reduced by 6%. The temperature in the POX unit was too high to promote the conversion of CO by the water gas-shift. In the third case, the POX unit increases the hydrogen yield from 19.4%v after pyrolysis to 31%v on a dry basis. The H 2 content in case 3 is lower than in case 2. First, a part of hydrogen is kept in bio-char. In addition, the carbon in biochar which is not converted into syngas as CO could not contribute to the H 2 formation by the water-gas shift reaction. To increase the amount of hydrogen after POX unit, we tried to add steam in the POX. This addition has no effect on H 2 formation. Indeed, as shown by our group [START_REF] Dufour | Mechanisms and Kinetics of Methane Thermal Conversion in a Syngas[END_REF] , OH radicals mainly result from CO 2 conversion during syngas thermal conversion and H 2 O is poorly reactive under such conditions.

Tar and particles

The Figure 6 shows the tar dew point and the tar concentration in the syngas along the process. The POX unit seems to have no effect on tar dew point. In fact, tars are effectively converted but the heaviest tars mainly control the dew point value even at very low concentration. A higher amount of oxygen would increase the efficiency of the POX unit. Unfortunately, the temperature rise would have been too high for the refractory material of the reactor. The addition of oxygen was then limited by the temperature reached in the POX unit (Figure 7) (more details presented in SI S5). The temperature limit was set to 1300°C for a classical refractory material of the POX unit. The peak temperature corresponds to the maximum temperature reached in the POX unit.

The reduction of global tar concentration is relatively small after the wet scrubber for case 1, but this equipment is required for water condensation from syngas. It also plays the role of extra dust/soot removal. The heaviest PAHs are removed by the wet scrubber, thus reducing the tar dew point.

In case 3, the POX unit is not able to reduce the amount of tars below 2 g/Nm 3 as for gasification scenarios, even with higher gas-phase residence times. The addition of oxygen was limited by the temperature reached (max 1300°C). The initial content of tars is higher after pyrolysis than after gasification. Its composition is also different with more primary and secondary tars. However, the tar dew point is reduced to around 25°C after the wet scrubber.

H 2 separation by a hybrid process: membrane coupled to PSA

Two stages of separation were required (membrane and PSA) when the H 2 concentration in the syngas was too low, since a single PSA would require a very large product recycling to reach 70%v at the inlet [START_REF] Spath | Biomass to Hydrogen Production Detailed Design and Economics Utilizing the Battelle Columbus Laboratory Indirectly-Heated Gasifier[END_REF] .

The membrane as a first stage plays the role of a pre-concentrator before the PSA (see supplementary material S7). To the best of our knowledge, this architecture of hydrogen separation from a biomass syngas is proposed for the first time. No data is available on the membrane lifetime using syngas with a residual amount of tars. If necessary, a guard bed filled with activated carbon could be added to remove these tars before the separation stage.

The minimum specific separation costs were estimated at 0.91 and 1.08 €/kg H2 for case 2 (Figure 8) and 3 respectively (SI S7). The first stage of separation with the membrane module requires a syngas pressure of around 5 bar, which is lower than the pressure required for PSA H 2 (around 25 bar). Only a part of the syngas enriched in hydrogen is compressed to this higher pressure in the second stage.

As a result, the required energy is lowered with these two levels of pressure. The operational expenditure is decreased for the power consumption, but the H 2 recovery rate is lower. This explains why the specific separation cost is only 0.59 €/kg H2 in the case 1 with one stage PSA.

Comparison of the various options and recommendations

The choice between these scenarios is dependent on various criteria. The heat represents roughly between one third and more than half of energy outputs. Therefore, the choice may depend on the valorization of heat on the selected locations (like industrial sites). In any case, it would be far better if the heat demand of the site is relatively constant during the year. The availability of the feedstock is another criterion on the scale that is linked to the location of the plant. Finally, this location should avoid the transport of hydrogen on large distance by trucks. The bio-char can be more easily transported than hydrogen, but its production supposes a demand nearby since its density is quite low. This bio-char could also be stored or sequestrated to create a carbon sink [START_REF] Dufour | Geological Sequestration of Biomass Char to Mitigate Climate Change[END_REF] .

Case 1 could fulfill hydrogen demand at 293 kg H2 /h and it produces 5.2 MW th of heat that must be valorized. Case 2 and 3 are simpler from a technology point of view: fewer unit operations, no catalytic reactor. This gain in robustness has to be counterbalanced with the hydrogen separation process that is more complex. The production of hydrogen is also far lower 98 kg H2 /h in case 2 and only 66 kg H2 /h in case 3.

From these results, we could expect the capital costs of case 1 to be much higher than cases 2 and 3 due to the catalytic reactors and the additional cost of catalysts. However, for cases 2 and 3, the hydrogen separation was achieved in two stages, increasing the specific cost of hydrogen separation.

This point should be quantified properly with a further techno-economic assessment.

The third case can be chosen for its carbon sequestration potential through bio-char production. This option leads to a net negative CO 2 emission process as it may lead to a stable sink of carbon [START_REF] Dufour | Geological Sequestration of Biomass Char to Mitigate Climate Change[END_REF] .

Furthermore, in the context of bioenergy with carbon capture and storage (BECCS) processes, the PSA tail gas has a high concentration of CO 2 (between 60 to 70%v according to the cases). This high concentration favors the capture of CO 2 for carbon sequestration.

CONCLUSION

The aim of this work was to provide a detailed mass and energy balance of three scenarios of production of hydrogen from biomass: a first case dedicated to produce the maximum of hydrogen, a simpler option without catalyst reactors and a carbon-negative process that also produced bio-char.

The model covered the whole process from drying of the biomass to the production of H 2 , heat and bio-char. All these operation units were modeled with Aspen Plus® with a detailed composition of tars.

The accuracy of the model was ensured with experimental data when they were available. Chemkin

Pro was coupled to Aspen Plus® to model the partial oxidation unit with a detailed radical kinetic mechanism. A hybrid hydrogen separation process was proposed using two technologies, namely membrane and PSA. The high-temperature heat was recovered for steam generation used for the gasification. The low-grade heat was used for woodchips drying whereas the rest of the heat was valorized in a heating network.

Global energetic efficiencies are 75.4, 77.8 and 80.4%net for scenarios 1 to 3, respectively. The hydrogen yields were 79, 26 and 18 g H 2 /kg biomass,dry after separation. The excess of heat dedicated to a heating network leads to heat efficiencies of 23.4, 60.0 and 49.0%net for the same three options respectively. 110 g biochar /kg biomass,dry was produced in the third option. The needs of utilities and commodities are also quantified.

These data will be used in a techno-economic assessment and a life cycle assessment to consider all the aspects: profitability and environmental impacts.

The production of renewable hydrogen from biomass represents an alternative path to electrolysis processes when the available electricity production is too low or too carbon intensive.

SUPPORTING INFORMATIONS

Figure 1 :

 1 Figure 1: Simplified process flow diagram and the various modeling approach for each units.

Figure 1

 1 Figure 1 presents a simplified flow sheet of the 3 cases and Figure 2 the detailed flow sheet for case 1.

Figure 2 :

 2 Figure 2: Detailed flow diagram of case 1.

Figure 3 :

 3 Figure 3: Sankey's type diagrams (in MW) of the 3 cases.

Figure 4 :

 4 Figure 4: Hydrogen molar flow rate and fraction along the process. G=Gasifier, P=Pyrolyzer, POX=Partial Oxidation, R=Reformer, WGS=Water Gas Shift, M=Membrane, PSA=Pressure Swing Adsorption.

  contains more tars than Corella et al. 11 with dolomite as bed material. Indeed the syngas produced by Corella et al. contains less than 2 g/Nm 3 of tars whereas, the syngas produced by Schmid et al. contains 38 g/Nm 3 17 .

Figure 5 :

 5 Figure 5: Fate of atomic hydrogen along the process for each scenario (kg H /h).

Figure 6 :

 6 Figure 6: Concentration of tars and tar dew point along the process. G=Gasifier, P=Pyrolyzer, POX=Partial Oxidation, R=Reformer, WS=Water Scrubber. Tars and benzene in blue, tars without benzene in green.

Figure 7 :

 7 Figure 7: Impact of ER on POX temperature and reduction of tars (for cases 1 and 2).

Figure 8 :

 8 Figure 8: Specific hydrogen separation cost for case 2 as a function of membrane area, permeate and retentate pressures.

Table 1 : Wet scrubber efficiency. Pilot plant, this work (Venturi + wet scrubber) a Rabou et al. (2009) 63

 1

	Water absorber

Table 2 : Assumptions of the Aspen Plus model, utilities and material consumption. 303

 2 Mass of fresh adsorbent per year estimated from Peters et al. and Swanson et al.[START_REF] Swanson | Techno-Economic Analysis of Biofuels Production Based on Gasification[END_REF][START_REF] Peters | Plant Design and Economics for Chemical Engineers[END_REF] , adsorption isotherms for lithium doped adsorbent[START_REF] Silva Ferreira | High-Purity Oxygen Production by VPSA[END_REF] assuming a 1-year lifetime (7500 hours of operations). d Assuming 2.6 kg/h of fresh bed material for a 20 MW gasifier.

	Unit		Methods		Results
	Drier b	Adapted from François et al. 27	Heat required, VOCs emissions
	VSA O 2 a,c	SEP, O 2 recovery rate 55%		Air input
	Gasifier d	RYIELD with external Fortran subroutine	Composition of the syngas,
		(experimental data 17 and atoms balances)	temperature reached
	Auto-thermal	RYIELD with external Fortran subroutine	Composition of the syngas and
	pyrolyser	(ChemkinPro)			bio-char, temperature reached
	Cyclone	SEP, ΔP = 0.5 kPa
	POX	Fortran	subroutine,	CHEMKIN-PRO	Composition of the syngas after
		simulation with adiabatic plug-flow reactor	POX unit, temperature reached
		and radical kinetic mechanism 57-59
	Steam reformer e RPLUG, kinetics in SI S6, T inlet 845°C	Composition and temperature of
						the syngas after reformer,
						amount of catalyst.
	Water-gas-shift f RGIBBS, temperature approach 20°C,	Composition and temperature of
		ΔP = 0.4 kPa, HTS (T inlet 350°C), LTS (T inlet	the syngas after WGS, amount of
		200°C)				catalyst.
	Water scrubber g FLASH	with	experimental	data,	Composition of the syngas after
		ΔP = 0.15 kPa water flow rate adjusted to	scrubber
		reach a syngas at 30°C at the outlet.
	Compressor a	Multi-stage compressor with intercooler	Power required and outlet
		(30°C), GPSA method, polytropic efficiency	temperature.
		0.80 and mechanical efficiency 0.98
	Membrane H 2 i	Cape-open MEMSIC 66 , countercurrent flow	Compositions of the outlets and
		pattern and permeance for UBE B-H	the corresponding membrane
		membrane from 67 .			surface area
	PSA H 2 h	SEP, recovery efficiency 85%, ΔP = 5 kPa	Flow of hydrogen produced,
						composition of the tail gas,
						amount of adsorbent.
	Air booster a	GPSA method, polytropic efficiency 0.80	Power required and outlet
		and mechanical efficiency 0.98	temperature.
	Gas boiler	RYIELD with external Fortran subroutine	Heat generated by the boiler,
		adapted from François et al. 27	exhaust gas composition.
		-excess of air λ=1.5
		-CO: 0.006 kg/Nm 3
		-C 10 H 8 : 4 10 -9 kg/Nm 3
		-Other PAHs: 1.5 10 -9 kg/Nm 3
		(acenaphthylene, anthracene,
		phenanthrene, pyrene)
		mole fraction 0.25 each.
		-NO mass fraction from oxidation of
		atmospheric N 2 7 10 -5
		-Soot 3 10 -5 kg/Nm 3
		-VOC 0.25 g/Nm 3
	Heat exchanger	ΔP = 2 kPa, minimum temperature	Surface area
		approach 5°C

a Estimated from Aspen Plus model assuming 80% polytropic efficiency and 98% mechanical efficiency for compressors and boosters. b Ref 20 is used for conveyor and dryer consumption.

c e

Table 3 :

 3 Utilities consumption.

				Case 1	Case 2	Case 3
	Electrical consumption [MWe]	2.08	1.66	1.36
	Fresh	water for	steam	25.9	9.4	0
	generation [thousands of m 3 /y]
	Adsorbent VSA O 2 [t/y]		0.96	0.96	0.63
	Bed material [t/y]		19.5	19.5	19.5
	Reformer catalyst [t/y]		0.25
	WGS catalyst [t/y]		HTS: 1.26
				LTS: 0.68
	Chemicals for water treatment NaOH: 210 t/y	NaOH: 210 t/y	NaOH: 210 t/y
				Flocculant: 23 m 3 /y	Flocculant: 23 m 3 /y	Flocculant: 23 m 3 /y
				Coagulant: 23 m 3 /y	Coagulant: 23 m 3 /y	Coagulant: 23 m 3 /y
				AC: 0.37 t/y	AC: 11.7 t/y	AC: 24.8 t/y
	Adsorbent PSA H 2 [t/y]		AC: 0.465	AC: 0.090	AC: 0.063
				Ze: 0.371	Ze: 0.071	Ze: 0.049
	Membrane H 2 area [m²/y]			300	300
	Natural gas a [Nm 3 /y]		3600	3600	3600
	Nitrogen b [Nm 3 /y]		3600	3600	3600

a Estimate for two start-ups per year and auxiliary fuel for flare. b Estimate for the nitrogen safety system.

  1 (79 g H 2 /kg biomass,dry ) is in good agreement with previous studies. In bubbling fluidized bed, Ersöz et al. evaluated a yield of 76.1 g H 2 /kg biomass,dry after PSA 36 . Susmozas et al. found a lower value in a steam dual fluidized bed: 55.0 g H 2 /kg biomass,dry 41 whereas Pallozzi et al. found 75.2 g H 2 /kg biomass,dry 42 for the same technology. Gupta and Dasappa determined 107.4 g H 2 /kg biomass,dry 37 with a fixed bed downdraft. The values obtained for case 2 and 3 of this

The supporting information file 1 presents: 1) details flowsheets of each scenarios and mass balance results, 2) the composition of the syngas and tars used as experimental data for oxy -steam gasification, 3) the presentation and the results of the auto-thermal pyrolyser model, 4) few details about the partial oxidation model, 5) the influence of the equivalent ratio on the efficiency of the partial oxidation unit, 6) the kinetic model used for the steam reformer model, 7) the methods and results used to define the architecture of the hydrogen separation. The supporting information file 2 presents the detailed composition of each flow from Aspen Plus.

FUNDINGS

This work was funded by the French PIA project "Lorraine Université d'Excellence" (reference ANR-15-IDEX-04-LUE) and by the Hy-C-GREEN project (Europe-FEDER and Grand-Est province).