Production of Hydrogen from Biomass with Negative CO₂ Emissions Using a Commercial-Scale Fluidized Bed Gasifier

Abstract

Biomass gasification, as a thermochemical method, has attracted interest due to the growing popularity of biofuel production using syngas or pure hydrogen. Additionally, this hydrogen production method, when integrated with CO₂ capture, may have negative CO₂ emissions, which makes this process competitive with electrolysis and coal gasification. This article presents the results of process and economic analyses of a hydrogen production system integrated with a commercial, fluidized-bed solid fuel gasification reactor (SES technology—Synthesis Energy Systems). With the use of a single gasification unit with a capacity of 60 t/h of raw biomass, the system produces between 72.5 and 78.4 t/d of hydrogen depending on the configuration considered. Additionally, assuming the CO₂ emission neutrality of biomass processing, the application of CO₂ capture leads to negative CO₂ emissions. This allows for obtaining additional revenue from the sale of CO₂ emission allowances, which can significantly reduce the costs of hydrogen production. In this analysis, the breakthrough price for CO₂ emissions, above which the hydrogen production costs are negative, is USD 240/t CO₂.

1. Introduction

As a result of increasingly stringent emission standards related to energy production, there has been a gradual shift away from traditional fossil fuels towards more diverse renewable sources and hydrogen. Hydrogen is currently one of the most desirable gases due to its wide range of applications in various industries [1,2]. The use of hydrogen enables the use and transport of energy from renewable sources on a much larger scale than just through direct electricity generation. It can also be used directly as a fuel that does generate CO₂ emissions and may allow for the diversification of domestic energy sources in the long term [3,4].

Currently, hydrogen is used as a fuel and a substrate in the production of methanol, ammonia, synthetic fuels and many others. Produced hydrogen is divided into four groups based on the type of raw material, the technology used and the emitted pollutants [5]: black hydrogen, which is hydrogen produced from coal through gasification; gray hydrogen, i.e., hydrogen produced from natural gas through its reforming (both of these processes are associated with CO₂ emissions); blue hydrogen, which is hydrogen produced from fossil fuels using CO₂ capture; and green hydrogen, i.e., emission-free hydrogen produced in the electrolysis process. Of course, these can be further separated by considering additional sources of electricity (energy from the grid, renewable energy or nuclear energy) [4,6]. Hydrogen production from biomass is also a potential source of green hydrogen, particularly when produced through the thermochemical conversion of biomass. The two main possible production pathways are biomass gasification and pyrolysis; however, the technological maturity of the pyrolysis process is currently assessed at a TRL (technology readiness level) of five [7]. This is mainly due to the high costs of raw material preparation and challenges related to scaling up the installations [7,8].

Currently, most commercial hydrogen production is associated with the use of the following fossil fuels: natural gas (62% of global production), coal (19%), oil (18%) and low CO₂ emission sources (only 1%). The potential of hydrogen is closely related to the reduction in CO₂ emissions from its production. Changing the current trends in hydrogen production still requires a large amount of research and financial resources that will allow the technology behind low-emission hydrogen production processes to mature, which is one of the goals of the net-zero 2050 program.

The use of electrolysis ensures the highest process efficiency, which can reach 75%; however, due to the very high demand for electrical energy, it is associated with the highest cost of production per kg of H₂. A second, more economically advantageous option is methods of producing hydrogen from biomass, among which are biological methods (fermentation, photolysis) and thermochemical methods (gasification, pyrolysis, liquefaction).

Two main trends in the development of hydrogen production technology can be observed in the literature: integration of hydrogen production with simultaneous generation of electrical or thermal energy, and improvements in process efficiency and reductions in the amount of emitted greenhouse gases.

The hydrogen yield of thermochemical biomass processing depends on many parameters, such as the biomass type, its grain size, the temperature, the catalysts used and the process time [9,10]. For this reason, there are many works in the literature devoted to both laboratory and simulation analyses of this process [11,12]. Ali Sabbaghi et al. [13] proposed a system allowing for the production of electrical and thermal energy using a thermodynamic process model. This system consists of a gas turbine system, an electrolyzer, a fuel cell and two CO₂ loops (TCO₂ and SCO₂). The obtained electrical energy production efficiency is 47.89% and 2.74 kg H₂/h, using a biomass stream of 198 kg/h. In addition, this study also examined the effect of using different types of biomass on the efficiency of the process. Hurtado et al. [14] studied gas yields for different biomass mixtures and the environmental safety of the process based on a developed thermodynamic model of biomass gasification. Li et al. [15] proposed a two-stage biomass gasification process, where the gasification reaction takes place at temperatures between 750 and 950 °C in the first stage, and the reforming process mainly takes place in the next reactor stage, achieving a hydrogen yield of 142 g/kg of biomass (dry and ash-free, daf). Frigo et al. [16] presented the results from a test run in a downdraft gasifier, where the hydrogen content in the produced syngas reached up to 40%, with a CGE (cold gas efficiency) of about 80%.

Fluidized bed reactors are also used in hydrogen production [17]. Erdem et al. [18], based on the Aspen Plus program, investigated the effect of the composition of the waste and biomass mixture and the steam-to-raw material ratio on the gas yield and the amount of hydrogen produced. Chen et al. [19] conducted similar studies using a pilot fluidized bed reactor (H = 60 cm) for the gasification of sewage sludge and its mixtures. Manu et al. conducted gasification of various types of biomass in a fluidized bed reactor (1–3 kg of biomass/h) using air as a gasification medium and using various biomass mixtures. It was found that the use of biomass with a higher content of alkali metal oxides increases the yield. Gil et al. [20] investigated the effect of an oxygen–steam mixture on the gasification of wood biomass, obtaining a maximum of a 30% hydrogen concentration in the process. Luo et al. [21] studied the effect of calcium oxide addition on the gasification of wood biomass in a fluidized bed reactor. In their work, they presented both a process model developed in the Aspen program and laboratory test results confirming that adding calcium oxide to the system increased hydrogen yields and reduced the release of liquid products. Biomass gasification, as a hydrogen production method, is a well-known and mature technology that allows the processing of various types of raw materials, including waste [22].

The cost of producing a kilogram of hydrogen depends on the technology and raw material used and changes with the scale of production [23]. As of 2021, over 80% of global hydrogen production came from fossil fuel technologies, which is largely due to their lower cost. However, as the cost of CO₂ emissions increases, the unit price of hydrogen production will approach that of currently more expensive, less emission-intensive technologies. A comparison of production costs and CO₂ emissions for the listed technologies is provided in Table 1.

Table 1. Comparison of hydrogen production technologies [5,22,24].

Color	Technology	Feed	TRL	LCOH 1 (USD/kgH2)	CO2 Emissions (kg/kg)
Grey	SMR 2	NG 3	9	0.7–2.5	8.5
	ATR 4	NG	7–9	1.2	11.0
	Coal gasification	Coal	9	1.9–2.5	22.0
Blue	SMR (CCUS) 5	NG	8–9	1.2–2.1	2.4
	ATR (CCUS)	NG	7–9	1.7	3.9
	Coal gasification (CCUS)	Coal	6–7	2.1–2.6	4.1
Turquoise	Methane pyrolysis	NG	3–4	2.2–3.4	0
Green	Biomass gasification 6 (with CCS)	Biomass	5	2.5–3.5	0
	Electrolysis	Water	6–9	2.6–8.5	0

¹ Levelized cost of hydrogen production. ² Steam methane reforming. ³ Natural gas. ⁴ Autothermal reforming. ⁵ Carbon capture, utilization and storage. ⁶ Some sources suggest that only hydrogen produced via electrolysis can be called green hydrogen [25].

In fluidized bed reactors, the gasification process takes place in a bed suspended in a stream of reaction gases. To ensure the correct course of the process, the solid fuel fed to the reactor must be crushed to a grain size of 0.5–6 mm. The residence time of the fuel in the reactor is usually in the range of 10–100 s. Fluidization ensures intensive contact of the solid and gas phases and, consequently, a uniform temperature distribution in the bed, which is maintained below the ash softening temperature (800–1050 °C). This is an extremely important requirement due to the possibility of slagging and, consequently, unstable reactor operation. Relatively low gasification temperatures are the cause of incomplete C conversion, which leads to a deterioration in the efficiency of the process and the possibility of impurities in the gas (tar substances). In order to increase the efficiency of the process, unreacted char is recirculated or char is burned in a separate combustion unit (hybrid system). The main advantage of fluidized bed reactors is their flexibility in terms of fuel quality and feed size. These reactors are particularly useful for converting low-calorie fuels, including those with a calorific value below 21 MJ/kg. However, they have a lower limit for effective fuel conversion into gas, defined by a calorific value of no less than 14 MJ/kg.

Fluidized bed technologies, along with the intensively developed entrained flow reactors, are the subject of increasing attention. Lower process temperatures than those of slagging reactors (below the ash melting point) reduce investment and operating costs and improve the reliability and availability of the system. Additionally, a high operating efficiency, moderate oxygen and steam demands and a high fuel flexibility make fluidized bed technologies an interesting alternative to gasification in entrained flow reactors.

The most technologically advanced solutions are the ones provided by KRW (Kellog–Rust–Westinghouse) [26,27,28,29,30,31], Uhde/HTW (High-Temperature Winkler) [28,32], U-Gas GTI/SESs (Synthesis Energy Systems) [33,34,35] and KBR Transport Reactors [36,37,38,39]. All provided technologies have been developed on a pilot or demonstration scale as well as for commercial applications (Table 2).

Table 2. Industrial fluid bed gasification reactors.

Location	Technical Parameters	Application	Year Started/Current Status
KRW (Kellog–Rust–Westinghouse) [27]
IGCC Pinion Pine Project, Reno, NV, USA	▪ 35 t/h, 1.9 MPa ▪ Air gasification	▪ IGCC demo plant (100 MW)	Started in 1996 Closed after demonstration period
Uhde/HTW (High-Temperature Winkler)
Rheinbraun AG, Berrenrath, Germany	▪ 25 t/h, 1.0 MPa ▪ Lignite coal	▪ Demo plant, methanol production	1986–1997
Rheinbraun AG, Wesseling, Germany	▪ 7 t/h, 2.5 MPa, ▪ Air/oxygen gasification	▪ Demo plant, future IGCC perspective plan for 300 MW IGCC	1989–1992
KBR
Power Systems Development Facility (PSDF), Wilsonville, AL, USA	▪ 2 t/h, 1.7 MPa, ▪ Air/oxygen gasification	▪ Pilot plant, Experimental runs for different gas utilization scenarios	1999
Stanton Energy Center, Orlando, FL, USA	-	▪ Commercial plant, IGCC, 285 MWe	Planned to start in 2010 No information about current project status
Tian Ming Electric Power in Dongguan, China	▪ 1600 t/day, lignite ▪ 1 reactor	▪ Commercial plant, IGCC	Planned to start in 2013 No information about current project status
Xilinguole Sunite Alkali Industry Co., Ltd., Inner Mongolia, China	▪ 1000 t/day, lignite ▪ 1 reactor	▪ Commercial plant, Production of chemicals	Planned to start in 2014 No information about current project status
Kemper County energy facility, Kemper County, MS, USA	▪ 8000 t/day (lignite ▪ 2 reactors	▪ Commercial plant, IGCC	Started in 2016 [40]
U–GAS GTI/SES [33,35,41]
Tampella Power inc., Tampere, Finland	▪ 30 t/day, 2.3 MPa (coal)	▪ Demo plant ▪ Electricity production	1992 In operation for 3800 h
Flex–Fuel Test Facility (FFTF), Des Plaines, IL, USA	▪ 10 t/d (air) ▪ 20 t/d (oxygen) ▪ 3.0 MPa	▪ Pilot plant ▪ Experimental runs for different gas utilization scenarios (liquid fuels, hydrogen, energy production)	2004
Shanghai Coking and Chemical Corporation, Wujin Chemical Industry Area, Changzhou, China	▪ 150 t/day, 8 reactors (2 in reserve) ▪ Air gasification ▪ Atmospheric pressure	▪ Commercial plant ▪ Syngas burned in a coking plant	Started in 1995, stopped and demolished due lack of syngas utilization
Shandong Haihua Coal & Chemical Co., Ltd., Zaozhuang, Shandong, China	▪ 28,000 mn3/h (phase I) ▪ 17,000 mn3/h (phase II)	▪ Commercial plant ▪ Methanol from syngas	In 2014, 2 reactors were in operation at 400 t/day each
YIMA Coal Industry Group Co., Ltd., Yima, Henan, China	▪ 1,000,000 t/year (methanol)	▪ Commercial plant ▪ Methanol from syngas	Started in 2016 (195,000 t of methanol was produced in 12.2016–09.2017)
Chalco Aluminum Corporation, Henan, Luoyang, China	▪ 4 reactors, 120,000 mn3/h of syngas	▪ Commercial plant	Started in March 2017
Chalco Aluminum Corporation Hejin, Shanxi, China	▪ 1 reactor 28,000 mn3/h syngas	▪ Commercial plant	Started in December 2015

SES Technology

One of the technologies using a fluidized gasification reactor is SESs. SGT (SES gasification technology) was created based on a modified U-Gas technology (Gas Technology Institute, Des Plaines, IL, USA). The differences compared to the U-Gas technology include the following:

▪

Fuel drying method;

▪

Use of coke to ignite the reactor, which reduces tar emissions during heating;

▪

Modified dosing system;

▪

Dust reduction system, modified cyclones and fuel recirculation system;

▪

Ash cooling system;

▪

Easier process scaling and increased production capacity and reactor pressure;

▪

Reduced CAPEX.

Thanks to the above changes, this technology is distinguished by its high flexibility regarding the range of raw materials that can be gasified, starting from coal with a low degree of metamorphism through to biomass and municipal waste. SGT was tested on a commercial scale using fuels of various origins and with different degrees of carbonization (Table 3) as follows:

Table 3. Characteristics of the fuel used in SGT reactors [41].

Coal Properties	Tested SGT Reactors
	ZZ, Zao Zhuang Plant	YIMA Plant	Experience of SGT
Ash (%mas.)	10–55	19–52	1–55
Moisture (%mas.)	4–43	1–10	1–43
Volatiles (%mas.)	12–40	20–30	3–69
Fixed carbon (%mas.)	24–66	22–38	6–83
Sulfur (%mas)	0.6–4.0	0.2–2.0	0.2–4.6
Ash melting point, °C	1112–1450	1277–1488	1040–1460
LHV, kJ/kg	12,950–25,500	10,950–21,250	10,950–32,200

▪

Hard coals—American, Polish, Chinese, Indian;

▪

Brown coals—Montana, Saskatchewan, Inner Mongolia;

▪

Coke and chars, including from tires;

▪

Biomass—willow, straw, wood pellets, rice husks, sorted municipal waste.

This technology can be used to produce synthesis gas for energy applications as well as for chemical synthesis (production of hydrogen, SNG—synthetic natural gas, chemicals or fertilizers). The simplicity of the SGT system results in it having a reliability and availability similar to or better than other solid fuel gasification technologies using a fluidized bed. SGT has less demanding conditions than high-temperature gasification technologies (GE, Shell, Siemens, CBI) and is mechanically simpler than other similar gasification technologies using a fluidized bed. SES technology has a high availability (>90%) and reliability (>98%). It is also characterized by a significant reduction in the content of condensable hydrocarbons in the raw gas compared to reactors operating at a similar temperature.

The key advantages of SGT in syngas production compared to other commercially available technologies include the following:

▪

Broad feedstock flexibility covering all coals, especially the lowest quality coals and coal waste, biomass and municipal waste;

▪

Ability to change and manage feedstocks during operation;

▪

Ability to produce energy from syngas at a significantly lower cost than from oil and LNG;

▪

Lower capital costs than previous generations of gasification systems;

▪

High cold gasification efficiencies;

▪

Lower efficiency loss when feeding low-quality coals;

▪

Significant amounts of methane in the syngas product;

▪

Lowest possible water consumption;

▪

Ability to modularize and build on both small and large scales.

The general configuration of the discussed plants is shown in Figure 1. The main parts of the setup include the following:

Figure 1. Block diagram of SES gasification technology.

▪

Feedstock preparation;

▪

A gasification island;

▪

A raw gas cleaning and cooling system;

▪

Molten ash and fly ash removal.

2. Materials and Methods

2.1. Goal of This Study

This paper presents both process and economic analyses of a hydrogen production system integrated with biomass fluidized bed gasification based on SES technology available on the market. The process model was validated using data from a large-scale commercially available SES plant. The economic analysis was performed with assumptions specific to the Polish energy market.

Simulations were run for the following three different plant configurations:

• Case A (base case): Hydrogen production with CO₂ capture. The steam reformer (that converts the methane contained in the raw gas) is fueled by natural gas supplied from outside the system (maximizing net hydrogen production).
• Case B: Hydrogen production with CO₂ capture. The steam reformer is fueled by a portion of the hydrogen produced within the system (leading to a reduction in net hydrogen production but a decrease in the CO₂ emissions related to methane combustion).
• Case C: Hydrogen production without CO₂ capture. The steam reformer is fueled by natural gas supplied from outside the system (maximizing net hydrogen production).

2.2. Computational Model of a Fluidized Bed Reactor—Synthesis Energy System (SES) Technology

SES gasification (formerly U-Gas) is oxygen/steam fluidized bed gasification used to produce synthesis gas for various applications. Five plants are now operating worldwide on an industrial scale, using 12 SES gasification reactors to produce methanol and gas for fuel (aluminum production plants). The developed model is based on a Gibbs reactor connected in series with a stochiometric reactor (Figure 2). Three additional reactions are assumed to occur in the stochiometric reactor to transform CO and N₂:

Figure 2. Model of the synthesis energy system (SES) gasification reactor—calculation scheme.

The CO shift reaction:

$$\mathrm{C}\mathrm{O}+{\mathrm{H}}_2\mathrm{O}\leftrightarrow {\mathrm{H}}_2+{\mathrm{C}\mathrm{O}}_2,-41 \mathrm{MJ}/\mathrm{kmol},$$

(1)

CO fractional conversion = 0.09 (Fractional conversion = (CO_inlet − CO_outlet)/CO_inlet) (Analogous for N₂ conversion)

The methane production reaction:

$$\mathrm{C}\mathrm{O}+{3\mathrm{H}}_2\leftrightarrow {\mathrm{C}\mathrm{H}}_4+{\mathrm{H}}_2\mathrm{O},-206 \mathrm{MJ}/\mathrm{kmol},$$

(2)

CO fractional conversion = 0.11

The ammonia production reaction:

$${\mathrm{N}}_2+{3\mathrm{H}}_2\leftrightarrow 2{\mathrm{N}\mathrm{H}}_3,-45.9 \mathrm{MJ}/\mathrm{kmol},$$

(3)

N₂ fractional conversion = 0.008.

The model was verified based on data obtained from an industrial gasification plant [42,43].

2.3. Enthalpy of Biomass Formation

The thermodynamic calculations of the gasification process included a calculation of the biomass enthalpy of formation. It has been shown in [43] that the effects of the enthalpy of formation can influence the thermodynamic results, including changes in temperature and the equilibrium composition of the produced syngas. The enthalpy of formation is defined according to Equation (4) [44]:

$${∆}_{\mathrm{f}}{\mathrm{H}}^0={\mathrm{Q}}_{\mathrm{s}}+{∆}_{\mathrm{c}}{\mathrm{H}}^0$$

(4)

The thermodynamic enthalpy of combustion is calculated based on the following formula:

$${∆}_{\mathrm{c}}{\mathrm{H}}^{0,\mathrm{}\mathrm{d}\mathrm{a}\mathrm{f}}=-327.633\mathrm{}{\mathrm{C}}^{\mathrm{d}\mathrm{a}\mathrm{f}}-1417.892\mathrm{}{\mathrm{H}}^{\mathrm{d}\mathrm{a}\mathrm{f}\mathrm{}}-92.768\mathrm{}{\mathrm{S}}_{\mathrm{c}}^{\mathrm{d}\mathrm{a}\mathrm{f}}, \mathrm{kJ}/\mathrm{kg}$$

(5)

while the calorimetric heat of combustion may be calculated using the following formula, depending on the amount of oxygen in the fuel (θ):

$${\mathrm{Q}}_{\mathrm{s}}=(-1){∆}_{\mathrm{c}}{\mathrm{H}}^0\mathrm{f}\left(\mathsf{\theta }\right)$$

(6)

2.4. Process Assumption for the Hydrogen Production System

An analysis of the possibility of producing hydrogen from biomass was carried out based on the results obtained from our own process model of the hydrogen production system integrated with biomass gasification developed in the ChemCAD process simulator. The developed model includes the basic technological nodes of the system with a defined calculation algorithm and adopted operating parameters, connected by process streams including gas streams, steam (at various pressure levels), water (process, cooling and boiler) and solids. The presented model is capable of determining the most important data on a hydrogen production plant, including the mass and energy balance, syngas composition, energy requirements and CO₂ emissions. The general process diagram of the system is shown in Figure 3. A hydrogen production system with a capacity of 60 t/h of raw biomass with a 40% moisture content (capacity of one gasification reactor) was assumed as the basis for calculations. The assumed availability is 7884 h, which gives a biomass flow rate of 473,040 t/year. The fuel characteristics are presented in Table 4.

Figure 3. Schematic diagram of hydrogen production plant from biomass gasification.

Table 4. Feedstock biomass properties.

Parameter	Symbol	Unit	As Received	Dry	Biomass Feed to the Reactor (20% Moisture)
Coal	C	%	31.42	52.37	41.89
Hydrogen	H	%	2.48	4.13	3.31
Nitrogen	N	%	0.13	0.22	0.17
Sulphur	St	%	0.01	0.02	0.01
Oxygen	O	%	24.26	40.43	32.35
Moisture	Wt	%	40.00	0.00	20.00
Ash	A	%	1.70	2.83	2.27
LHV	Wd	kJ/kg	9742	17,903	13,824
HHV	QS	kJ/kg	11,300	18,833	15,069

The analyzed H₂ production system consists of the following elements:

▪

An oxygen and nitrogen compression system;

▪

A biomass preparation system;

▪

A biomass gasification system with gas cooling and purification;

▪

A CO conversion system;

▪

A gas cooling system;

▪

A carbon dioxide removal system;

▪

A H₂ separation system;

▪

A char and residual gas combustion system (from the PSA installation, steam and exhaust gas production to the biomass drying system).

Specific solutions for individual technological nodes were selected considering the following aspects: technological suitability due to the fuel characteristics (including the moisture content in the biomass), the production strategy, the availability of the technology on a commercial scale and the process efficiency.

2.4.1. Oxygen and Nitrogen Compression System

Oxygen and nitrogen for the gasification system were supplied from an external supplier. It was assumed that the temperature and pressure of oxygen and nitrogen at the boundary of the plot are 32 °C/8.5 bar and 32 °C/4 bar, respectively (typical parameters of oxygen and nitrogen produced in a high-pressure air separation system) [45]. Before being fed to the gasification installation, the oxygen and nitrogen streams are compressed to a pressure of 40 bar in a compressor system with interstage cooling.

2.4.2. Fuel Preparation System

The system consists of a biomass crushing and drying unit for the reactor. Biomass from the raw material storage facility is directed by a stacker–loader and a belt conveyor system for screening to a vibrating screen, where it is separated into fine and coarse fractions. The coarse fraction from the screen is transported by a belt conveyor system to the crusher, while the fine fraction from the vibrating screen and the resulting crushed coarse fraction are directed by a belt conveyor system and bucket conveyors to the drying unit—a drum dryer system. Heat for the drying process is supplied diaphragmatically in the form of hot exhaust gases resulting from the combustion of char and residual gases [45,46]. The partially dried fuel (moisture content: 20%) is directed to buffer tanks, while vapors and exhaust gases from the dryers are released into the atmosphere after dust removal. Drying the biomass to a moisture content of 20% ensures that the lower calorific value of the fuel is at an acceptable level of 14 MJ/kg, which is necessary for efficient fuel processing in fluidized bed reactors. Figure 4 illustrates the relationship between the cold gasification efficiency and the combustion heat of the raw gas relative to the moisture content of the fuel supplied to the reactor in the analyzed case.

Figure 4. The impact of biomass moisture content on cold gasification efficiency and the heat of combustion of raw gas after the gasifier.

2.4.3. Gasification System

The gasification system consists of a fuel dosing system, a gasification reactor and a gas pre-cleaning and cooling system. Fuel from the drying unit’s buffer tanks is directed to the raw material tank via belt conveyors, a circular bucket conveyor and a screw conveyor. The gasification reactor is supplied by three feeding systems with a maximum efficiency of 50% each. During normal operation, each raw material feeding system operates at 1/3 of the reactor’s efficiency. Fuel is fed to the reactor by a screw feeder and then a pneumatic feeder, with three nozzles feeding fuel to the fluidized reactor. The pneumatic feeder is supplied with compressed nitrogen or CO₂. The gasification reactor is a fluidized bed reactor according to the SES technology (previously U-Gas) and operates at a pressure of 30 bar. Oxygen and steam are used as gasifying and fluidizing agents and are introduced into the lower part of the reactor.

Hot process gas is discharged from the reactor to the cyclone, where solid particles are separated; the remainder is separated in the second-stage cyclone. The separated char and ash particles are returned to the gasification reactor to maximize the conversion of C in the gasification process. Uncleaned, hot synthesis gas (approx. 1000 °C) leaving the secondary cyclone is fed to the steam reforming unit (SMR), where methane and aliphatic and aromatic hydrocarbons (BTX, oils, tar) contained in the raw gas from biomass gasification are decomposed. The heat for the process is supplied by burning natural gas (Cases A and C) or produced hydrogen (Case B). For calculation purposes, the SMR process temperature was assumed to be 995 °C. The gas leaving the reforming unit flows to the synthesis gas cooler to recover heat. In the gas coolers, medium- and low-pressure superheated steam (60, 18 bar) is produced. Part of the medium-pressure steam (approx. 60%) is fed to the gasification system, and the excess is used to produce electricity in the steam turbine system. The gas, cooled to approx. 180 °C, is directed to a high-efficiency cyclone and then to a water scrubber to remove halides, phenols, NH₃, HCN, residual fine solids and the rest of heavier hydrocarbons unconverted in the SMR unit. Water from the bottom of the scrubber is recirculated to the top of the column. In order to avoid accumulation of the removed pollutants, approx. 10% of the water is discharged from the system to the purification station (wastewater). Water separated from the gas in the coolers before the CO₂ removal system is used to replace this water.

2.4.4. Process Gas Cleaning and Conversion

After leaving the gasification unit, the raw gas is subjected to CO conversion (CO-shift), CO₂ cooling and separation processes, and then hydrogen release. The CO conversion process is an extremely important element of the system as it is a significant source of hydrogen (for the analyzed case, approximately 37% of the hydrogen produced comes from the CO conversion process). The gas supplied to the conversion unit contains the required moisture content and there is no need to supply water vapor to it (for effective conversion, the molar ratio of water vapor to CO in the gas before the CO-shift unit should be >2). Low process temperatures are conducive to achieving a high degree of CO conversion, but at a low reaction rate. For this reason, a two-stage process is most often used to intensify the process (high- and low-temperature conversion with interstage cooling) [45,47,48]. For the analysis, it was assumed that the system is equipped with two-stage CO conversion equipment. After the first conversion stage, the gas is cooled to a temperature of about 200 °C and fed to the low-temperature conversion reactor (stage II). During gas cooling, superheated, medium-pressure steam (60 bar) is produced, which is used to produce electricity in a steam turbine. Shifted gas is directed to deep desulfurization in a ZnFe₂O₄ bed [49]. The hydrogen production process also requires carbon dioxide removal in order to obtain a gas stream with a high H₂ concentration before the PSA unit. As a result, a stream of carbon dioxide with a high mass concentration (usually above 99%) is obtained as a by-product, which can be sold or transported and stored. At the typical pressures leaving the gasification reactor, the most economical separation method from the point of view of energy demand is physical absorption. A CO₂ absorption process, carried out using Selexol technology (assumed separation efficiency of 92%), was adopted for further analysis. In order to achieve an appropriate CO₂ separation efficiency, it is necessary to cool the gas to about 35 °C before CO₂ removal. The gas is cooled in three stages: low-pressure steam (4.5 bar) is produced, the water supplied to the system is heated (to 100 °C) and water coolant (supplied from outside the system) is used.

2.4.5. Hydrogen Separation System

For hydrogen separation, PSA (pressure swing adsorption) technology is assumed. The PSA process is commonly used in industry for large-scale hydrogen production, and it allows high-purity hydrogen to be obtained with a separation efficiency of 80–95% [48,50,51]. Taking into account the technology’s maturity, the potential for large-scale application and the process conditions—such as the composition and pressure of the gas produced—the PSA process is currently the most favorable technological option. Table 5 presents a comparison of hydrogen recovery systems utilizing PSA technology, membrane separation and cryogenic methods. A hydrogen separation efficiency (in the PSA system) of 95% was assumed for the calculations.

Table 5. Characteristics of the separation technology [50].

Parameter	PSA	Membrane Separation	Cryogenic Separation
Basic process performance indicators
Minimum H2 concentration at inlet	50	15	15
Inlet gas pressure [MPa]	1.03–6.9	1.38–13.8	1.38–8.3
Outlet H2 concentration [%]	>99.9	98	97
H2 separation/recovery efficiency [%]	<95	<97	<98
CO and CO2 separation capacity	Yes	No	No
Outlet pressure (H2) relative to inlet	Comparable	Significantly lower	Comparable
General operational characteristics
Feed pretreatment	No	Yes	Yes
Flexibility to inlet gas composition changes	Very high	High	Average
Reliability	High	High	Average
By-product recovery	No	Possible	Yes
Easy to expand	Average	High	Low

2.4.6. CO₂ Separation Unit

Carbon dioxide leaving the separation system (Selexol) at two pressures (1.2 and 5.5 bar) is compressed to a pressure of 100 bar in a compressor with interstage cooling [45]. Liquefied CO₂ with parameters of 10 MPa/20 °C is prepared for transport to potential storage locations.

2.4.7. Steam Turbine

The superheated steam produced in the system at pressures of 60, 18 and 4.5 bar is used to produce electricity in a condensing steam turbine. The steam parameters at the inlets to the high-pressure, medium-pressure and low-pressure parts of the turbine are, respectively, 60 bar/460 °C, 17 bar/290 °C and 4 bar/150 °C. The turbine system generates power at the level of 13 MWe.

2.4.8. Auxiliaries

The analyzed system also includes a boiler water compression system and a combustion boiler fueled by the char from the gasification process and tail gas from the PSA system. The exhaust gases from the boiler are used to dry the biomass and for additional steam production (60 and 4.5 bar), which, together with the steam produced during cooling of the process gas (gasification, CO-shift and gas cooling system before the Selexol process), is used to produce electricity.

A summary of the system configuration data is shown in Table 6.

Table 6. Plant specification and assumptions in the base case scenario.

Specification	Unit	Values
Oxygen production	–	▪ N2 and O2 are produced outside of the considered model ▪ Inlet parameters: O2, 32 °C/8.5 bar; N2, 32 °C/4 bar ▪ Oxygen purity: 95% vol.
Gas compression unit	-	Three-step oxygen compressor and four-step nitrogen compressor
Gasification island
Reactor	–	SES fluidized bed
Pressure	bar	30
Ratio O2/C	kg O2/kg C	▪ 0.85 ▪ 0.36 with respect to biomass stream 1
Ratio steam/C 1	kg H2O/kg C	▪ 1.2 ▪ 0.5 with respect to as-received biomass 1
Carbon conversion	%	95.0
Syngas LHV 2	kJ/mn3	3960 (calculated)
CGE (cold gas efficiency)	%	71 2
Gas cooling	–	Convective cooling, produced steam at 60.18 bar
Gasifying agent	–	▪ Oxygen (95% vol.) ▪ Steam 60 bar (produced in the system)
Fuel	–	Dried biomass (20% m. moisture content)
Fuel transport gas	-	Nitrogen, 7500 kg/h
Initial gas processing	-	Cyclones, water scrubber
Gas conditioning and WGS reactor
WGS (CO conversion)	–	Two-step reactor, CO conversion: 96% (calculated)
Desulphurization	–	Adsorption ZnFeO4
Claus process	–	Not included
CO2 separation	–	Selexol, effectiveness: 92%
Hydrogen separation
Technology	–	Pressure swing adsorption (PSA)
Efficiency PSA	%	95
Biomass drying
Heat source	-	Combustion of char and tail gas
Final moisture content	% m.	20%
Additional units
Energy production	-	Steam turbine, 60/17/4 bar
Cooling water	°C	15/26 °C
CO2 compression	-	▪ Seven-step compression unit with interstage cooling ▪ Outlet CO2 parameters: 100 bar/20 °C

¹ Biomass feed to the reactor (20% m. moisture content). ² Raw gas after steam reforming.

2.5. Production Cost Calculation—Methodology

The basic economic parameters used to estimate the cost of producing H₂ and electricity are given in Table 7. This table contains data on all costs related to the operation of the installation, and the values presented were determined based on market price analyses and source materials.

Table 7. Basic economic assumptions.

Parameter	Value	Comment
Biomass price	30 PLN/GJ	Own assumption according to market price
Capacity factor	90%	Yearly availability
Oxygen price	84.3 USD/t	Own estimation according to commercial data, Poland, 250 PLN/t
Nitrogen price	84.3 USD/t	Own estimation according to commercial data, Poland, 250 PLN/t
Electricity price	304.63 PLN/MWh	Monthly market electricity price, Polish market, December 2023 [52]
Carbon tax	77.98 EUR/t CO2	CO2 Emissions Futures Contracts, December 2023 [53]
USD and EUR valued in a year	2023, December	rate EUR/PLN: 4.348 rate USD/PLN: 3.935 Average rates NBP (1), December 2023 [54]
Operation and maintenance costs	5.3 to 5.6% of total investment cost (see Table 5)	Calculated separately for each case, including chemicals, water, waste disposal, insurance, tax, renewals
Transport and storage costs	13.59 USD/t	Calculated using data acquired for year 2019 adjusted using PPI for transportation and warehousing industries up to December 2023 [55]

⁽¹⁾ NBP: Central Bank of the Republic of Poland.

To calculate the CAPEX, an exponential investment assessment and price growth index methods were used.

The exponential investment assessment method is based on the following function:

$${\mathrm{C}}_1={\mathrm{C}}_0{\left({\displaystyle \frac{{\mathrm{S}}_1}{{\mathrm{S}}_0}}\right)}^{\mathrm{f}}$$

(7)

where C₁ is the calculated investment for the system component, C₀ is the reference investment cost, S₁ is the scale of the system component, S₀ is the base scale parameter and f is the scaling exponent.

The base scales and scaling exponents for the components of the production facilities based on data from the literature and in-house experience are shown in Table 8.

Table 8. Parameters used to estimate purchased equipment cost.

No.	Plant Component	Scaling Parameter S0	S0 Value	Unit	Scale Factor f	Co [mln USD] 1	Base Year	Source
1	Biomass processing and drying	Raw biomass feed	83,333	kg/h	0.75	3.840	2005	[56]
2	Gasification island (gasifier, syngas cooler, preliminary gas cleaning, fines/ash handling)	Therma input (fuel)	258	MWth	0.67	18.498	2017	[41]
3	N2 compression	Compression power	10	MWe	0.67	1.260 2	2002	[57]
4	O2 compression	Compression power	10	MWe	0.67	1.690 2	2002	[57]
5	CO2 drying and compression	Compression power	13	MWe	0.67	3.968 2	2002	[57]
6	CO shift	Gas stream	161	Mg/h	0.67	0.866	2005	[56]
7	Sulphur removal, adsorption	Sour gas stream	81	Mg/h	0.67	1.204	2005	[56]
8	Selexol (CO2 removal)	Pure CO2 captured	327	Mg/h	0.67	8.800 2	2002	[57]
9	PSA H2	H2 separated	6468	kg/h	0.6	5.041	2005	[56]
10	Off gas boiler	Therma input (fuel)	236	MWth	0.67	8.222	2007	[51]
11	BFWP 3	Water stream	158,425	kg/h	0.33	0.096	2005	[56]
12	Steam turbine	ST gross power	136	MWe	0.67	15.871 2	2002	[57]
13	CH4 reformer	Methane stream	5845	kg/h	0.7	4.966	2005	[56]

¹ Equipment purchase cost. ² Cost of equipment calculated from overnight capital costs using Peters and Timmerhaus factors [56]. ³ Boiler feed water pump.

2.5.1. Price Growth Index Method

The next step of determining the cost of specific equipment in the plant is converting its value from the base year to the current year of implementation. The price growth index method used is characterized by the following equation:

$${\mathrm{C}}_2={\mathrm{C}}_1\left({\displaystyle \frac{{\mathrm{I}}_2}{{\mathrm{I}}_1}}\right)$$

(8)

where C₂ is the calculated investment cost for the current year, C₁ is the reference investment cost for the base year, I₂ is the price growth index value for current year and I₁ is the price growth index value for the base year.

Capital expenditures specified for the base year were calculated for the current year using CEPCIs (chemical engineering plant cost indices, Figure 5). To calculate total investment costs (TICs), Peters and Timmerhaus factors were used according to Table 9.

Figure 5. Price growth index for chemical plants—CEPCI.

Table 9. Factors for total investment cost estimation [56].

Investment Cost Component	% of TPEC
Total purchased equipment cost (TPEC)	100.0
Purchased equipment installation	39
Instrumentation and controls	26
Piping	31
Electrical systems	10
Buildings (including services)	29
Yard improvements	12
Total installed cost (TIC)	247.0
Indirect costs
Engineering	32
Construction	34
Legal and contractors’ fees	23
Project contingency	37
Total indirect costs	126.0
Total project investment	373.0

2.5.2. Levelized Cost of Hydrogen Production (LCOH) Calculation Methodology

LCOH calculations were performed based on the LCOE (levelized cost of energy) calculation methodology according to the Cost Estimation Methodology for NETL Assessments of Power Plant Performance [58].

The following equation was used to calculate the LCOH index value. All calculations were performed using the nominal approach:

$$\mathrm{L}\mathrm{C}\mathrm{O}\mathrm{H}=\mathrm{L}\mathrm{C}\mathrm{C}+\mathrm{L}\mathrm{O}\mathrm{M}+\mathrm{L}\mathrm{F}\mathrm{P}$$

(9)

where:

• LCC—levelized cost of capital;
• LOM—levelized annual O&M expenses;
• LFP—levelized annual fuel expenses.
• LCC and LOM were obtained according to Equations (10)–(15):

$$\mathrm{L}\mathrm{C}\mathrm{C}=\mathrm{T}\mathrm{I}\mathrm{C}\times \mathrm{F}\mathrm{C}\mathrm{R}$$

(10)

$$\mathrm{F}\mathrm{C}\mathrm{R}={\displaystyle \frac{\mathrm{C}\mathrm{R}\mathrm{F}}{(1-\mathrm{E}\mathrm{T}\mathrm{R})}}-{\displaystyle \frac{\mathrm{E}\mathrm{T}\mathrm{R}\times \mathrm{D}}{(1-\mathrm{E}\mathrm{T}\mathrm{R})}}$$

(11)

$$\mathrm{C}\mathrm{R}\mathrm{F}={\displaystyle \frac{\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C}\times {(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C})}^{\mathrm{y}}}{{(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C})}^{\mathrm{y}}-1}}$$

(12)

$$\mathrm{D}=\mathrm{C}\mathrm{R}\mathrm{F}\times {\sum }_{\mathrm{n}=1}^{\mathrm{z}}{\displaystyle \frac{{\mathrm{d}}_{\mathrm{n}}}{{(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C})}^{\mathrm{n}}}}$$

(13)

$$\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C}=\%\mathrm{E}\mathrm{q}\mathrm{u}\mathrm{i}\mathrm{t}\mathrm{y}\times {\mathrm{r}}_{\mathrm{E}\mathrm{q}\mathrm{u}\mathrm{i}\mathrm{t}\mathrm{y}}+\%\mathrm{D}\mathrm{e}\mathrm{b}\mathrm{t}\times {\mathrm{r}}_{\mathrm{D}\mathrm{e}\mathrm{b}\mathrm{t}}\times (1-\mathrm{E}\mathrm{T}\mathrm{R})$$

(14)

$$\mathrm{L}\mathrm{O}\mathrm{M}=\mathrm{A}\mathrm{O}\mathrm{M}\times {\displaystyle \frac{\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C}\times {(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C})}^{\mathrm{y}}}{{(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C})}^{\mathrm{y}}-1}}\times {\displaystyle \frac{1-{\left[\frac{\left(1+\mathrm{i}\right)}{\left(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C}\right)}\right]}^{\mathrm{y}}}{\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C}-\mathrm{i}}}$$

(15)

where:

• TIC—total investment cost;
• FCR—fixed charge rate;
• CRF—capital recovery factor;
• ETR—effective tax rate;
• ATWACC—nominal after tax weighted average cost of capital;
• D—current value of tax depreciation expense;
• d_n—tax depreciation fraction in year n;
• y—number of operating years;
• z—number of years of depreciation;
• AOM—annual O&M costs;
• i—assumed annual O&M escalation rate.

In the case of determining the LFP value, for the available price paths, the source material [58] provides the following equation:

$$\mathrm{L}\mathrm{F}\mathrm{P}={\mathrm{P}\mathrm{V}}_{\mathrm{f}\mathrm{u}\mathrm{e}\mathrm{l}\mathrm{}\mathrm{p}\mathrm{r}\mathrm{i}\mathrm{c}\mathrm{e}}\times {\displaystyle \frac{\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C}\times {(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C})}^{\mathrm{y}}}{{(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C})}^{\mathrm{y}}-1}}$$

(16)

where PV_fuelprice equals:

$${\mathrm{P}\mathrm{V}}_{\mathrm{f}\mathrm{u}\mathrm{e}\mathrm{l}\mathrm{}\mathrm{p}\mathrm{r}\mathrm{i}\mathrm{c}\mathrm{e}}=\sum _{\mathrm{n}=1}^{\mathrm{y}}{\displaystyle \frac{{\mathrm{P}}_{\mathrm{n}}}{{(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C})}^{\mathrm{n}}}}$$

(17)

However, in the analyzed cases, it was assumed that the fuel price updates would be calculated following the same method as for O&M costs. Therefore, an alternative equation was adopted, analogous to Equation (15):

$$\mathrm{L}\mathrm{F}\mathrm{P}=\mathrm{F}\mathrm{P}\times {\displaystyle \frac{\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C}\times {(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C})}^{\mathrm{y}}}{{(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C})}^{\mathrm{y}}-1}}\times {\displaystyle \frac{1-{\left[\frac{\left(1+\mathrm{i}\right)}{\left(1+\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C}\right)}\right]}^{\mathrm{y}}}{\mathrm{A}\mathrm{T}\mathrm{W}\mathrm{A}\mathrm{C}\mathrm{C}-\mathrm{i}}}$$

(18)

The numerical assumptions adopted for the calculations and the calculated unit indicators of investment, operating and fuel costs are summarized in Table 10.

Table 10. Basic assumptions for LCOH calculation methodology.

Parameter	Value	Notes
%Equity	50%	Assumed share of equity and debt for the analyzed project
%Debt	50%
requity	10%	Assumed following NETL assumptions [58]
rdebt	8%	Assuming average value of WIBOR 1 6 M factor 6% + margin 2%
ETR	19%	Average CIT 2 value
dn	5%	Average straight-line depreciation rate for period of 20 years
z	20
y	30	Assumed following NETL assumptions [58]
i	2.5%	The long-term average annual dynamics of the CPI index were adopted in accordance with the guidelines for the use of uniform macroeconomic indicators [59]

¹ WIBOR—Warsaw Interbank Offer Rate—polish indicator representing interest rate on loans between banks, usually defines base value for variable interest rate on loans. ² CIT—Corporate Income Tax.

3. Results and Discussion

3.1. Gasification Reactor

The mass and energy balance of the gasification system is presented in Table 11. The system consists of a gasification reactor, raw gas cooling (production of steam at 60 and 18 bar) and a gas scrubber.

Table 11. Mass and energy balance in the gasification reactor.

Stream	t [°C]	p [bar]	Mass Flow [kg/h]	Chemical Enthalpy [MW]	Physical Enthalpy [MW]	Total [MW]
Biomass 1	43.4	1.0	45,141.0	188.3	1.4	189.7
Oxygen (95% vol.)	47.4	40.0	17,000.0	0.0	0.1	0.1
Transport gas (nitrogen)	56.7	40.0	7500.0	0.0	0.1	0.1
Natural gas (steam reformer)	25.0	2.0	665.2	10.3	0.0	10.3
BFW, 60 bar	100.0	60.0	36,946.1	0.0	3.3	3.3
BFW, 18 bar	100.0	18.0	5700.1	0.0	0.5	0.5
Air	25.0	1.2	13,671.4	0.0	0.0	0.0
Process water (scrubber)	54.8	35.0	2800.0	0.0	0.1	0.1
Total input	-	-	129,423.8	198.6	5.3	203.9
Process gas	176.0	27.0	92,144.6	141.1	22.5	163.6
Steam, 60 bar	540.0	59.7	14,324.1	-	13.6	13.6
Steam, 18 bar	226.2	17.7	5700.1	-	4.4	4.4
Char	25.0	1.2	1962.6	8.6	0.0	8.6
Flue gases	1000.0	1.2	14,336.9	-	5.6	5.6
Waste water (from scrubber)	164.4	27.0	955.9	-	0.2	0.2
Losses 2	-	-	-	-	-	8.1
Total output			129,423.8	149.7	46.2	203.9

¹ Moisture content: 40%. ² Heat losses (reactor: 2% of the fuel chemical enthalpy; heat loss due to char cooling and other processes).

From the gasification of 45.1 t/h of semi-dried biomass, the reactor produces 90.3 t/h of raw gas. This gas, with hydrogen and methane contents of 21% vol. (1699 kg/h) and 2.4% vol. (1558 kg/h), respectively, is directed to the SMR (steam methane reforming) unit, where methane is decomposed and an additional 478 kg/h of hydrogen is produced. The system also produces 37 and 5.7 t/h of steam at 60 and 18 bar, respectively. After meeting the gasification reactor’s steam requirements, the net production of steam at 60 bar is 14.3 t/h. The solid product leaving the installation from the gasification process (char) contains 48%wt. of elemental carbon (the assumed conversion degree of elemental carbon in the gasification reactor is 98%) and has a calorific value of 15.7 MJ/kg. This product is used to generate heat to dry the biomass. After scrubbing, the gas (92.1 t/h) is fed into the WGS (water gas shift) system, where additional hydrogen is produced in the CO conversion process. The efficiency of the cold gasification is 71% (HHV). The obtained results are in good agreement with the data in [60,61,62] and slightly worse than those in [16]. The characteristic parameters of raw gas (at the outlet from the gasification system, after the scrubber) are presented in Table 12.

Table 12. Raw gas characteristics.

Parameter	Unit	Value
Gas yield	kg/h	92,145
Temperature	°C	176
Pressure	bar	27
LHV (HHV)	MJ/kg	4.26 (5.51)
Gas composition	% vol.
CO		14.8
CO2		19.0
H2		24.8
H2O		34.1
N2		6.4
CH4		0.4
Ar		0.4

3.2. Hydrogen Production Plant

Figure 6 contains a schematic diagram of the production plant, as well as the calculated values of the most important streams for the base case scenario.

Figure 6. Diagram of the hydrogen production plant with biomass gasification showing the characteristics of the main process streams (Case A).

The gasification reactor is unchanged in all three scenarios; hence, its cold gas efficiency remains stable at 71%. The hydrogen production efficiency, defined as the ratio of hydrogen chemical enthalpy to feedstock chemical enthalpy, is equal to 65% for Cases A and C and decreases to 49.6% for Case B. The relatively high efficiency of hydrogen production results from the assumed high efficiency of the PSA system (95%). The impact of the hydrogen separation efficiency in the PSA installation on the efficiency of the entire production system is important. For the lowest assumed PSA efficiency (75%), the hydrogen production efficiency reaches 51.3%, while after increasing the PSA efficiency to 95%, the hydrogen production increases up to 64.9%.

Table 13 presents a summary of the data from the literature on the efficiency of hydrogen production in systems integrated with solid fuel gasification equipment.

Table 13. Comparison of hydrogen production efficiencies for hydrogen production systems integrated with solid fuel gasification.

No.	Feedstock	Gasification Technology	Hydrogen Production Efficiency	PSA Removal Efficiency	Ref.
1	Coal, bituminous	GE Texaco—radiant cooler (Air Product)	60.81 (HHV)	0.8	[51]
2	Coal, bituminous	GE Texaco—quench cooler (Air Product)	61.05 (HHV)	0.8	[51]
3	Biomass	Battelle Columbus Laboratory (BCL) low-pressure, indirectly heated gasifier	55.5 (HHV)	0.8	[56]
4	Biomass	Downdraft, fixed bed, low pressure	63.8 (LHV) 1 68.01 (HHV) 1	0.95	[63]
5	Biomass	Dual fluidized bed (DFB)	68.9 (LHV)	n.a.	[64]
6	Coal/biomass (co-gasification)	Shell—convective cooler/quench (Air Product)	57.7 (HHV)	0.85	[65]
7	Coal, bituminous	Shell—convective cooler/quench (Air Product)	55.4 (LHV)	0.85	[66]
8	Lignite	Shell—convective cooler/quench (Air Product)	56.3 (LHV)	0.85	[66]

¹ Calculated assuming an SMR efficiency of 90% and a CO conversion (WGS) of 98%.

The total amount of produced hydrogen is 3268 kg/h for Cases A and C and 3022 kg/h for Case B. Although Case B considers methane reforming to increase hydrogen production, the heat needed to perform the reaction is produced via hydrogen combustion, resulting in a total decrease in process efficiency. The total production of hydrogen decreases by 7.6% (246 kg/h) compared to reforming using natural gas to produce the necessary energy. The total hydrogen production is 54.47 kg H₂/kg biomass for Cases A and C and 50.37 kg H₂/kg biomass for Case B.

A summary of the calculation results obtained for the three analyzed cases is presented in Table 14.

As a result of measures to decrease greenhouse gas emissions, newly developed technologies are required to stay within strict limits regarding produced CO₂. For this reason, all simulated scenarios were analyzed to calculate their CO₂ emissions and the possibility of zero or negative emissions for hydrogen production. The following study includes different sources of emitted CO₂, such as emissions from fuel combustion for the SMR unit, the total in situ emissions and additional emissions connected to energy production to supply electricity.

The simplest way to limit CO₂ emissions is by installing an additional carbon capture unit, which was simulated in Cases A and B. This results in a significant drop in the amount of CO₂ released to the atmosphere, from 70,894 kg/h to 12,037 kg/h and 10,212 kg/h, respectively, for Cases A and B.

However, due to the use of a CO₂ separation system, mainly the CO₂ capture and compression unit, the process’ electricity consumption increases to 77.4 kWh/Mg CO₂. As a result, the demand for electricity increases from 8.60 MWh (Case C) to approx. 13.60 MWh (Cases A and B), which, when converted to per kg of hydrogen produced, is 1.53 kWh and 1.65 kWh, respectively. Moreover, in Cases A and B, the electricity produced directly in the installation does not cover the system’s own needs, which is associated with the need to purchase additional power. If this energy does not come from renewable energy sources, it will be linked to additional, indirect CO₂ emissions. If biomass processing is considered CO₂ neutral, the specific CO₂ emissions in Cases A and B due to CO₂ capture become negative (−16.87 and −18.70 kg CO₂/ kg H₂, respectively). The CO₂ emission loads of the hydrogen produced in each of the three analyzed cases are significantly lower than those for hydrogen production from fossil fuels, both with and without CO₂ separation. Figure 7 presents a comparison of the emission intensity of hydrogen production for technologies using fossil fuels (coal gasification and thermal conversion of coke oven gas [66]), natural gas steam reforming [67]).

Figure 7. Comparison of carbon emissions for different hydrogen production technologies (own calculations [66], data from the literature [67]). The lower ranges correspond to the sequestration of CO₂ produced in hydrogen production systems.

Table 14. Main process parameters and emissions from biomass gasification hydrogen production plants with different configurations.

No	Parameter	Unit	Case A w/CCS 1	Case B w/CCS_H2 2	Case C w/o CCS
1	Biomass
	▪ Flow	kg/h	60,000.00	60,000.00	60,000.00
	▪ Chemical enthalpy	MWth	188.33	188.33	188.33 3
2	Methane
	▪ Flow rate	kg/h	665.19	0.00	665.19
	▪ Chemical enthalpy	MWth	10.26	0.00	10.26 3
2	Cold gas efficiency 4	%	71.06	71.06	71.06
3	Electric energy
	▪ Nominal power input	MW	11.22	10.61	10.53
	▪ Power requirements	MW	13.63	13.60	8.60
	▪ Net production 5	GWh	−18.95	−23.51	15.18
	▪ Energy usage	kWh/kg H2	4.17	4.50	2.63
4	Water/steam requirements ▪ BFW ▪ Cooling water	kg/h kg/h	124,159 30,129	125,287 27,999	126,103 22,169
5	Hydrogen production
	▪ Net production	kg/h	3268.37	3022.29	3268.37
		t/y	25,767.86	23,827.71	25,767.86
	▪ Hydrogen to biomass	kg H2/Mg biomass	54.47	50.37	54.47
	▪ Hydrogen efficiency 6	%	65	49.6	65
6	CO2 emission
	▪ CO2 emissions from SMR unit	kg/h	1824.79	0.00	1824.79
	▪ Total in situ CO2 emissions (hydrogen production)	kg/h	12,037.43	10,212.59	70,894.39
	▪ CO2 emissions due to energy consumption 7	kg/h	1894.28	2358.09	−1517.53
	▪ Total yearly emissions	Mg/y	109,289.75	80,516.05	573,318.04
	▪ Total yearly emissions including energy production	Mg/y	124,224.24	99,107.19	561,353.83
	▪ Negative emissions	kg/h	55,144.25	56,505.23	0.00
		t/y	434,757.24	445,487.21	0.00
7	Emission factors ▪ Emissions from SMR unit	kg CO2/kg H2	0.56	0.00	0.56
	▪ Total emissions from the plant assuming zero CO2 emissions from biomass processing	kg CO2/kg H2	−16.87	−18.70	0.094
8	Carbon capture efficiency 8	%	85.16	85.16	-

¹ The CO₂ released in the production process is compressed and prepared for transport. ² The source of heat in the SMR process is the combustion of part of the hydrogen produced. ³ Heat of combustion—HHV. ⁴ Ratio of the chemical enthalpy of the process gas (after the gasification reactor and reforming unit) to biomass (heat of combustion). ⁵ Annual production (plant availability: 90%). ⁶ Ratio of the chemical enthalpy of hydrogen to the chemical enthalpy of biomass (heat of combustion). ⁷ 788 kg CO₂/MWh [68]. ⁸ Efficiency of CO₂ removal in relation to the carbon contents in the fuel (efficiency of CO₂ removal from the process gas—92%).

Table 14. Main process parameters and emissions from biomass gasification hydrogen production plants with different configurations.

No	Parameter	Unit	Case A w/CCS 1	Case B w/CCS_H2 2	Case C w/o CCS
1	Biomass
	▪ Flow	kg/h	60,000.00	60,000.00	60,000.00
	▪ Chemical enthalpy	MWth	188.33	188.33	188.33 3
2	Methane
	▪ Flow rate	kg/h	665.19	0.00	665.19
	▪ Chemical enthalpy	MWth	10.26	0.00	10.26 3
2	Cold gas efficiency 4	%	71.06	71.06	71.06
3	Electric energy
	▪ Nominal power input	MW	11.22	10.61	10.53
	▪ Power requirements	MW	13.63	13.60	8.60
	▪ Net production 5	GWh	−18.95	−23.51	15.18
	▪ Energy usage	kWh/kg H2	4.17	4.50	2.63
4	Water/steam requirements ▪ BFW ▪ Cooling water	kg/h kg/h	124,159 30,129	125,287 27,999	126,103 22,169
5	Hydrogen production
	▪ Net production	kg/h	3268.37	3022.29	3268.37
		t/y	25,767.86	23,827.71	25,767.86
	▪ Hydrogen to biomass	kg H2/Mg biomass	54.47	50.37	54.47
	▪ Hydrogen efficiency 6	%	65	49.6	65
6	CO2 emission
	▪ CO2 emissions from SMR unit	kg/h	1824.79	0.00	1824.79
	▪ Total in situ CO2 emissions (hydrogen production)	kg/h	12,037.43	10,212.59	70,894.39
	▪ CO2 emissions due to energy consumption 7	kg/h	1894.28	2358.09	−1517.53
	▪ Total yearly emissions	Mg/y	109,289.75	80,516.05	573,318.04
	▪ Total yearly emissions including energy production	Mg/y	124,224.24	99,107.19	561,353.83
	▪ Negative emissions	kg/h	55,144.25	56,505.23	0.00
		t/y	434,757.24	445,487.21	0.00
7	Emission factors ▪ Emissions from SMR unit	kg CO2/kg H2	0.56	0.00	0.56
	▪ Total emissions from the plant assuming zero CO2 emissions from biomass processing	kg CO2/kg H2	−16.87	−18.70	0.094
8	Carbon capture efficiency 8	%	85.16	85.16	-

¹ The CO₂ released in the production process is compressed and prepared for transport. ² The source of heat in the SMR process is the combustion of part of the hydrogen produced. ³ Heat of combustion—HHV. ⁴ Ratio of the chemical enthalpy of the process gas (after the gasification reactor and reforming unit) to biomass (heat of combustion). ⁵ Annual production (plant availability: 90%). ⁶ Ratio of the chemical enthalpy of hydrogen to the chemical enthalpy of biomass (heat of combustion). ⁷ 788 kg CO₂/MWh [68]. ⁸ Efficiency of CO₂ removal in relation to the carbon contents in the fuel (efficiency of CO₂ removal from the process gas—92%).

3.3. Results of Economic Calculations

As part of this work, an economic analysis was carried out, including the determination of investment outlays and operating costs (CAPEX, OPEX) and the estimation of the costs of hydrogen production (levelized cost of hydrogen production, LCOH).

Calculations were carried out in accordance with the methodology and assumptions described in Section 2, the results of simulations of the analyzed systems (material and energy balances, characteristics of the main process streams) and investment cost estimates. The results of the calculations are presented in Table 15.

Table 15. Economic calculation results for biomass gasification hydrogen production plants with different configurations.

Parameter	Case A w. CCS	Case B w. CCS_H2	Case C w/o CCS
	mln USD 2023 (USD/kWHHv)	mln USD, 2023 (USD/kWHHv)	mln USD, 2023 (USD/kWHHv)
CAPEX	221.29 (1716.53)	220.65 (1850.89)	187.40 (1453.60)
OPEX total	43.34 (336.17)	37.41 (313.80)	74.61 (578.72)
▪ Raw materials	65.03 (504.40)	60.60 (508.34)	65.03 (504.40)
- Biomass	44.32 (343.80)	44.32 (371.80)	44.32 (343.80)
- Natural Gas	4.43 (34.34)	-	4.43 (34.34)
- Nitrogen and Oxygen	16.28 (126.26)	16.28 (136.54)	16.28 (126.26)
▪ Media, Energy, media, chemicals, T&S and environmental costs	−25.16 (−195.15)	−26.64 (−223.47)	1.33 (10.32)
- CO2 emission allowances	−37.46 (−290.57)	−38.39 (−321.99)	-
- others	10.80 (83.75)	11.75 (98.52)	1.33 (10.32)
▪ Fixed costs (including O&M costs)	9.78 (75.83)	9.76 (81.83)	8.65 (67.11)
LCOH	3.33	3.28	4.43

Under the adopted assumptions, the hydrogen production cost (LCOH) without CO₂ separation was USD 4.43/kg H₂ (Case C). Implementing CO₂ capture (Cases A and B) and accounting for revenues from the sale of CO₂ emission allowances (due to negative CO₂ emissions) reduced the production costs to approximately USD 3.3/kg H₂, representing a 25% decrease. Among the CO₂ capture variants, slightly lower production costs were observed in Case B, despite its lower hydrogen production efficiency compared to Case A (49.6% vs. 65%, Table 15). This was primarily due to the reduced operating costs caused by initiatives such as eliminating the need for natural gas and increasing negative CO₂ emissions, which boosted revenue from selling CO₂ allowances. The obtained results are in good agreement with the cost presented by Castro et al. [69], who calculated the cost of hydrogen from biomass as 3.71 USD/kgH₂. Excluding CO₂ transportation and storage (T&S) costs further lowered hydrogen production costs to USD 3.02 and USD 2.94/kg H₂ for Cases A and B, respectively. The production cost structure for the analyzed cases (taking into account T&S costs) is presented in Figure 8.

Figure 8. Hydrogen production cost structure for the cases considered.

The costs of hydrogen production depend on many factors, including production technology and fuel and energy prices, which can vary significantly depending on the location of the system. The obtained hydrogen production costs, for the adopted assumptions, are competitive with the current (as of 2022) costs of production from fossil fuels with CO₂ capture, which are around USD 1.9–7.2/kg H₂ and USD 2.1–5/kg H₂ for natural gas and coal, respectively, and significantly lower than hydrogen produced in the electrolysis process (renewable electricity), the production costs of which have been estimated at USD 3.4–12/ kg [1]. Despite the forecasted, significantly reduced costs of hydrogen production in 2030 (net zero emission scenario) both using fossil fuels (~1–3 USD/kg H₂) and electrolysis (~2–8 USD/kg H₂), hydrogen production from biomass still appears economically attractive. The undoubted advantage of using biomass for hydrogen production is the possibility of generating so-called negative emissions, which leads to negative hydrogen production costs [70]. The simulation shows that for the analyzed base case, the LCOH becomes negative CO₂ emission allowances above USD 240/kg H₂ (Figure 9).

Figure 9. The impact of the price of CO₂ emission allowances on the cost of hydrogen production.

Figure 10 presents the sensitivity analysis of the LCOH to changes in investment expenses, biomass purchase prices and CO₂ emission allowances. The analysis was performed in the range of −30%-+50%, and showed that the LCOH value is most affected by changes in biomass costs as well as CO₂ costs. A change in the biomass purchase price in the range of −30 to + 50% (~66–140 USD/Mg wet biomass) affects the change in LCOH in the range of 2.67 to 4.43 USD/kg H₂. Similarly, a change in the CO₂ emission allowance price from 60 to about 130 USD/Mg CO₂ (−30/+50%) causes a decrease in the LCOH from 3.9 to 2.4 USD/kg H₂. With the projected increase in emission allowances to 130 EUR/Mg CO₂ in 2040 [71] and assuming a constant biomass purchase cost, the cost of hydrogen production would drop to ~2 USD/kg H₂.

Figure 10. The impact of biomass purchase costs, CO₂ emission allowance prices and investment outlays on the percentage change (A) and absolute change (B) in LCOH.

4. Conclusions

This article presents designs and analysis results of a hydrogen production system integrated with commercial fluidized bed gasification technology offered by Synthesis Energy Systems. The analysis was carried out for a gasification reactor with a capacity of 60 Mg/h of raw biomass (40% moisture). With such a fuel consumption, depending on the adopted configuration, the system produces between 72.5 and 78.4 t/d of hydrogen.

Under the assumptions regarding the scale of this system, the costs of hydrogen production range from 3.28 to 4.43 USD/kg H₂, which means that this technology is currently competitive with both electrolytic methods (using renewable energy) and methods using fossil fuels (SMR, gasification).

Hydrogen production integrated with biomass gasification and CO₂ capture is the only technology that can achieve negative CO₂ emissions. This allows for zero or even negative hydrogen production costs for the established-limit cost of CO₂ emissions, which for the analyzed case (Case A) is USD 240/t CO₂.

The successful implementation of a hydrogen production system integrated with biomass gasification and CO₂ capture currently hinges on two main factors: the availability of raw materials and the development of infrastructure for CO₂ transport and geological storage. Wood biomass is considered the most suitable feedstock for gasification processes. Poland has significant potential in this area, with estimated resources of 13–16 million cubic meters per year from forest waste and the wood industry [72]. However, legislative and logistical challenges could impede the development of large-scale systems.

While biomass gasification and CO₂ separation technologies are well advanced, effectively integrating these processes into a hydrogen production system requires robust infrastructure for CO₂ transport and storage. Building and operating this infrastructure poses substantial technical and organizational challenges, necessitating expertise and practical experience. To address this, it will be crucial to launch pilot and demonstration projects focused on CCS, along with conducting long-term research during their operation. Globally, there are currently 628 projects related to carbon capture and storage (CCS) technology at various stages of progress. Of these, 50 are operational, 44 are under construction, and the remainder are in design or conceptual phases [73,74]. In Europe, notable projects include Norway’s Northern Lights [73,74], which plans to inject 3 million tons of CO₂ per year into the North Sea, and a Polish–Lithuanian initiative aimed at developing CO₂ transport infrastructure with a capacity of around 3 million tons annually by 2025–2030 [73,75].