Syngas Chemical Looping Process: Design and Construction of a 25

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Syngas Chemical Looping Process: Design and Construction of a 25 kWth Subpilot Unit Deepak Sridhar, Andrew Tong, Hyung Kim, Liang Zeng, Fanxing Li,† and Liang-Shih Fan* William. G. Lowrie Department of Chemical and Biomolecular Engineering 140 W 19th Avenue, 125 A Koffolt Laboratories, The Ohio State University, Columbus, Ohio 43210 ABSTRACT: The syngas chemical looping (SCL) process employing the gas−solid counter-current flow pattern demonstrates an innovative approach to generate hydrogen and/or electricity from syngas accompanied with in situ carbon capture. Iron-based oxygen carriers donate oxygen for complete syngas conversion in the reducer. The reduced oxygen carriers are then oxidized by steam and/or air to generate hydrogen and/or heat in the oxidizer and/or the combustor, respectively. Previous studies have reported the performance of the iron-based oxygen carriers, the advantages of a moving bed reducer and oxidizer, and simulation of various parametric effects on the reactor design of the reducer, oxidizer, and combustor for a continuous system. In this study, a 25 kWth subpilot SCL unit was designed based on the simulated criteria and constructed to demonstrate the feasibility of generating high purity hydrogen with in situ carbon capture. Two test runs were presented using 4.5 mm × 2.5−4.5 mm cylindrical oxygen carriers comprising of 60 wt % iron oxide (Fe2O3). The first test resulted in a syngas conversion of 99.96% and trace amounts of hydrogen generation highlighting the importance of the extent of oxygen carrier conversion. The second test demonstrated the continuous production of hydrogen with an average purity of 94.4% and a maximum of 98.4% when the conversion of the oxygen carriers exiting the reducer was 35.54%. The initial test results support the concept of continuous hydrogen generation with in situ carbon capture using the SCL process and also highlight the advantage of adopting the countercurrent moving bed reactor design.

1. INTRODUCTION A fast growth in hydrogen production is forecasted due to increased demand for power and chemicals such as ammonia, methanol, olefins, and transportation fuels.1 Commercially, hydrogen is generated from nonrenewable solid, liquid, and gaseous fossil energies, as their renewable counterparts (solar, wind, and biomass) are less economical.2 Steam−methane reforming is used to produce hydrogen from natural gas as shown in reaction a. CH4 + H 2O → CO + 3H 2 (a)

CO2 emissions are a growing concern in today’s society. Research at both lab and commercial scales is being performed to validate carbon capture, storage, utilization, and sequestration techniques.5 Precombustion carbon capture processes have been extensively researched, as they provide higher process efficiency compared to postcombustion carbon capture. Carbon capture is conventionally performed in precombustion processes by first shifting the syngas to CO2 and H2 via the water−gas shift reaction, as shown in reaction d:

This process is relatively efficient compared to traditional hydrogen generation processes from coal. However, steam− methane reforming can only be economically practical in areas where natural gas is an inexpensive feedstock. Gasification of coal and other carbonaceous fuels is another means employed to generate the hydrogen intermediate.3 Gasification refers to the partial oxidation of the fuel source to a gaseous fuel stream that retains majority of the original heating value of the feedstock. Syngas produced from gasification is expected to increase by 70% worldwide before 2015 where 45 GWth of coal will be processed for power generation via the integrated gasification combined cycle (IGCC).4 IGCC generates work from gas combustion and is more efficient compared to traditional coal boilers.3 The fuel gas stream produced from coal gasification is predominantly CO and H2, as depicted in reactions b and c, similar to steam−methane reforming.

Afterward, liquid absorbers or solid sorbents are used to remove the acid gas from the flue stream.6 However, such processes require either temperature or pressure swings resulting in significant process energy penalties and may be susceptible to fouling from syngas contaminants. Using these CO2 capture schemes, the cost of hydrogen production increases by 3−5% and 10−15% from natural gas reforming and coal gasification, respectively.7 The chemical looping concept has been studied as an alternative method for hydrogen production. The steam−iron process was first proposed in the early 20th century to utilize the iron-based particles for hydrogen production.8,9 This process performs oxidation−reduction cycles between metallic iron and Fe3O4 with steam and CO (or H2) as the feedstock. Reactions e and f given below summarize the reduction and oxidation reactions, respectively.

Cx Hy + H 2O → xCO + (y + 2)/2H 2

(b)

Cx Hy + x /2O2 → xCO + y/2H 2

(c)

© 2012 American Chemical Society

CO + H 2O → CO2 + H 2

(d)

Received: December 30, 2011 Revised: March 14, 2012 Published: March 15, 2012 2292

dx.doi.org/10.1021/ef202039y | Energy Fuels 2012, 26, 2292−2302

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Fe3O4 + CO (or H 2) → FeO/Fe + CO2 (or H 2O)

(e)

FeO/Fe + H 2O → H 2 + Fe3O4

(f)

sizing. The construction of the unit was performed based on the design criteria and operating conditions. The continuous operation of the subpilot SCL process demonstrates the capability of the chemical looping scheme to cogenerate concentrated CO2 and H2. The continuous operation of subpilot unit is an important proof of concept study to determine the feasibility of discrete CO2 and H2 cogeneration. To date, extensive prototype demonstrations have been conducted on chemical looping combustion processes for power generation with carbon capture.21,37−43 However, most chemical looping processes for hydrogen generation are only conceptual at this time. The subpilot SCL unit is the first demonstration unit for hydrogen production with 100% CO2 capture using iron-based oxygen carrier particles. 2. Design Criteria for the Subpilot SCL Unit. The subpilot SCL unit is the first demonstration of a continuously operating chemical looping combustion/gasification system embedding a moving bed reactor design concept. The SCL process uniquely uses two counter-current moving bed reactors allowing for a higher iron-oxide based oxygen carrier conversion compared to fluidized bed, while achieving complete gaseous fuel conversion. The higher oxygen carrier conversion provides the opportunity to generate H2 at a very high purity. Bench scale studies revealed that more than 99.5% purity of H2 can be generated using the moving bed reactor design, while achieving 100% syngas conversion.32,44 The SCL subpilot unit can be operated under two modes, namely, combustion mode and hydrogen generation mode. The combustion mode uses a counter-current moving bed reducer and fixed bed combustor. The hydrogen generation mode introduces a counter-current moving bed oxidizer between the reducer and the combustor. The oxidizer serves to generate pure H2 using steam oxidation of the iron oxide-based oxygen carriers. It is clear that the mode of operation governs the design specification of the subpilot unit. The processing capacity of the unit is defined as the amount of fuel that can be handled. The fuel for the SCL process is syngas derived from a coal gasifier. Therefore, it is important to account for the gasifier thermal efficiency to determine the capacity of the unit. The subpilot unit is operated using simulated syngas from gas cylinders. The overall syngas flow (Qsyn), in standard liter per minute (SLPM), has its main constituents of carbon monoxide and hydrogen where the volumetric flow rates are given by QCO and QH2, respectively, in SLPM. The thermal efficiency of the gasifier is denoted by ηg. The capacity of the unit (C), in kWth, can be arrived using the following equation.

The steam−iron process conducted had low fuel−gas conversion capability and particle redox recyclability.8,9 The oxidation−reduction reaction cycling study using several metal oxide based oxygen carriers, including iron oxide, gained impetus since the 1980s for in situ CO2 separation using the chemical looping combustion process.10,11 The chemical looping combustion process, mainly studied with methane as the fuel, can be performed using many transition metal oxides as the active component in the oxygen carrier particles.12−16 Specifically, nickel and copper based oxygen carriers have been extensively studied for chemical looping combustion application.17−30 Except iron, the other active metal components used in the oxygen carriers such as Ni, Cu, and Mn are thermodynamically restricted from producing H2 from steam economically as a result of low steam conversions.31−35 Coupling the merits of the steam−iron method and the chemical looping combustion process and introducing a unique reactor design, the iron-based syngas chemical looping (SCL) process was developed for the cogeneration of heat and hydrogen with in situ CO2 capture . The SCL process uses syngas from coal gasifier and utilizes iron-based oxygen carrier particles in a redox cycle. Concentrated hydrogen and CO2 are generated in discrete gas streams due to the reaction path design. This process scheme has the advantage of generating the CO2 stream at system pressure and temperature, reducing the energy penalty associated with carbon capture. Additionally, the CO2 removal and hydrogen/heat generation are integrated into one process. The SCL system consists of three main components: the reducer, oxidizer, and combustor. In the reducer, The Fe2O3 particles are reduced by the syngas feed to a mixture of metallic iron and wustite, while the syngas is converted CO2 and steam. The reduced particles are discharged from the reducer into the oxidizer. Here, the steam−iron reaction takes place producing hydrogen and partially oxidized Fe3O4 particles. The particles then travel to the combustor where air is used to fully oxidize the solids to Fe2O3 and convey them back to the reducer, completing the cycle. The combustion step is an exothermic reaction where the heat generated can be recovered to compensate for any parasitic energy requirements in the system as well as for power generation. Below is a list of the predominant reaction occurring in the three process components.32,36 Further details on the SCL process for commercial heat and material integration and process configurations can be found in previous literature.32,36,47 Fe2O3 + 2CO + 3H 2 (Syngas) → 2CO2 + 3H 2O + 2Fe (Reducer) Fe +

4 4 1 H 2O → H 2 + Fe3O4 (Oxidizer) 3 3 3

Fe3O4 +

C= (g)

⎡ Q CO 28 HHVCO ⎤ + ⎡ Q H2 2 HHVH2 ⎤ ⎢⎣ 22.4 60 ⎥⎦ ⎣ 22.4 60 ⎦ 1000ηg

(1)

2.1. Reducer Design. The counter-current moving bed reactor design needs to incorporate several factors required for a smooth solids flow. The superficial gas velocity (Usg) inside the reducer must be maintained below the minimum fluidization velocity of the particles (Umf). Additionally, for smooth solid flow inside the reactor, it is common practice to size the reactor diameter (ϕR) at least greater than ten times the oxygen carrier particle diameter (ϕP) to enhance solid flow and avoid solid plugging. The reducer is designed keeping the above two criteria in mind.

(h)

1 3 O2 (Air) → Fe2O3 + Heat (Combustor) 4 2 (i)

In this study, the design, construction, and continuous operation of a subpilot SCL unit were performed. The design criteria for the reactor system is first discussed where the particle performance results are used to determine the reactor 2293


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relationship between the height (Hred) and the diameter of the reactor (ϕ red).

The syngas entering the reducer comprises mainly of a combination of CO and H2. The gases counter-currently interact with the oxygen carriers to completely convert to a mixture of CO2 and steam, while reducing the particles. The oxygen carriers exiting the reducer have donated oxygen for fuel combustion before moving to the next reactor. This reaction scheme necessitates a reducer design that provides an adequate residence time for both the oxygen carrier and fuel to maximize their conversions. The reducer should completely convert the fuel and, therefore, provide adequate oxygen from the oxygen carriers to the fuel for complete combustion reactions: 1 CO + O2 → CO2 (j) 2

Vred = π

1 O2 → H 2O (k) 2 The reactions indicate that every mole of CO/H2 requires 1 /2 mole of oxygen molecule or 1 oxygen atom. This can be used to determine the oxygen demand (Θred), in g/min, required for complete fuel conversion by using the equation below:

Θred

mox − mfr mox − mr

χoxi = oc

1 1 1 H 2O → Fe3O4 + H 2 3 3 3

(k)

(l)

mox − mfo mox − mr

(9)

The oxygen demand in the oxidizer (Θoxi), in g/min, can be calculated using eq 10. The equation indicates that the conversion in the reducer needs to be greater that 0.11 to generate H2 using the chemical looping process. This is possible using the counter-current moving bed contacting pattern in the reducer.

(3)

red oxi Θoxi = [χ OC − χOC ]ṁ OCΨOC

(10)

The only reactive species introduced to the oxidizer is steam. Therefore, the steam flow requirement (QH2O), in SLPM, can be calculated from the oxygen demand in the oxidizer and the steam conversion (χoxi g ).

(4)

(5)

⎫ ⎧⎡ Θ ⎤ Q H O = ⎨⎢ oxi ⎥22.4⎬/χoxi 2 ⎭ g ⎩⎣ 16 ⎦

The moving bed reactor is similar to the ideal plug flow reactor. This means that the oxygen carrier residence time (τred OC), in minutes, required for maximum solid conversion is related to the volume of the reactor (Vred), the particle bulk density (ρb), in g/cm3, and the oxygen carrier flow rate. ⎡ ṁ ⎤ red τOC = VRed /⎢ OC ⎥ ⎢⎣ ρb ⎥⎦

(8)

The oxygen carriers undergo a transformation from Fe/FeO to Fe3O4. The oxygen carrier conversion in the oxidizer (χoxi OC) plays a crucial role in the determining the oxygen demand in the oxidizer. Equation 9 calculates χoxi OC, and it should be noted that the ideal value for χoxi OC is 0.11, as this represents complete conversion to Fe3O4. Values greater than 0.11 indicate incomplete oxidation of the oxygen carriers in the oxidizer.

Based on mass balance, the oxygen carrier flow rate (ṁ OC), in g/min, in the reducer can be derived using eqs 2−4. red ṁ OC = Θred /[χ OC ΨOC]

⎧ (ϕred)2 ⎫ ⎪ ⎪ 101325T ⎬ Usgred = Q syn /⎨ π ⎪ 4 ⎪ ⎩ ⎭ 298.15P

FeO +

(2)

The oxygen carrying capacity (ΨOC) of the oxygen carrier is defined as the maximum oxygen that can be donated by the oxygen carrier. This capacity accounts for the active metal oxide loading (δ) in the oxygen carrier and the oxygen mass ratio in the active metal oxide (Rmax).

ΨOC = δR max

(7)

Fe + H 2O → FeO + H 2

The oxygen demand in the reducer is provided by the oxygen carrier. The oxygen carrier flow rate can be determined using the oxygen demand and knowledge of the oxygen carrier conversion and carrying capacity. The oxygen carrier conversion (χred OC) is defined as the ratio of the amount of oxygen donated for fuel conversion to the maximum amount of oxygen that could be donated by the active metal oxide in the oxygen carrier. Here, mox, mrf , mr represent the mass of the oxygen carrier sample in the fully oxidized state, after reaction with the reducing fuel, and in its fully oxygen depleted state, respectively. red χ OC =

HRed

4

The superficial gas velocity can be calculated using the eq 8. Based on eqs 1−8 and the knowledge of the oxygen carrier properties, careful selection of the height and diameter of the reducer can be made allowing for smooth operating conditions coupled with optimum performance of the reducer. 2.2. Oxidizer Design. The oxidizer operates as a moving bed reactor similar to the reducer. This operating mode is designed to maximize steam oxidation of the reduced oxygen carriers to the magnetite phase coupled with high steam conversions based on the following reactions. Complete oxidation to Fe2O3 is not possible because of thermodynamic limitations.44

H2 +

⎡ Q CO + Q H ⎤ 1 2 ⎥ 32 =⎢ 22.4 ⎣ ⎦2

(ϕred)2

(11)

To prevent any accumulation of solids in the system, the oxygen carrier flow rate in the oxidizer is the same as that in the reducer. Therefore, the relationship between the residence time of the oxygen carriers in the oxidizer and the reactor volume can be established by eq 12 given below.

(6)

⎡ ṁ ⎤ oxi τOC = Voxi /⎢ OC ⎥ ⎢⎣ ρb ⎥⎦

Assuming a cylindrical reducer reactor helps determine the dimension of the reactor. The reactor dimensions provide the 2294

(12)


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manner similar to that adopted for the reducer and oxidizer, for a continuous flow combustor. The subpilot combustor, on the contrary, is designed to be operated in a batch fixed bed mode. Therefore, it is designed to hold a solid inventory for a given residence time to completely oxidize the oxygen carriers. The volume of the combustor (Vcom) can be determined from the solid circulation rate (ṁ OC) and residence time required in the combustor (τcom OC ) for complete oxidation.

Similar to the reducer, a cylindrical oxidizer reactor sets the reactor height (Hoxi) and diameter (ϕoxi) dimensions using eq 13. Voxi = π

(ϕoxi)2 4

Hoxi

⎧ (ϕ )2 ⎫ ⎪ ⎪ 101325T Usgoxi = Q H O/⎨ π oxi ⎬ ⎪ 2 4 ⎪ ⎩ ⎭ 298.15P

(13)

⎡ ṁ ⎤ com Vcom = ⎢ OC ⎥τOC ⎢⎣ ρb ⎥⎦

(14)

The superficial gas velocity can be calculated using eq 14, and a suitable reactor dimension can be determined to avoid oxygen carrier fluidization. 2.3. Combustor Design. The SCL process has the flexibility to be adapted for hydrogen generation, heat generation, or cogeneration. The final application of the SCL unit will determine the design of the combustor. On a larger scale, the combustor is expected to be a dense fluidized/ entrained bed reactor design. Fluidized bed combustors using iron-based oxygen carriers have been demonstrated in other chemical looping combustion demonstrations. The combustor serves a dual purpose in the SCL unit, complete solids oxidation, and conveyance to reducer. The combustor completely oxidizes the oxygen carriers to their initial state generating heat from the exothermic reactions. The oxidation reactions are shown below. 1 3 Fe3O4 + O2 → Fe2O3 (m) 2 4 FeO +

1 1 O2 → Fe3O4 6 3

Vcom = π

Q air ∈ 100

(1 + )

Hcom

(18)

orH 2 = ΘΔcom

22.4 32·0.2095

⎧ (ϕcom)2 ⎫ ⎪ ⎪ 101325T ⎬ Usgcom = Q air /⎨ π ⎪ 4 ⎪ ⎭ 298.15P ⎩

(19)

(20)

The air flow rate is maintained in the combustor for a predetermined residence time of the oxygen carriers in the combustor, and then the entrainment air is initiated. The entrainment air is introduced at a high flow rate to entrain the particles back to the reducer and complete the loop. Therefore, the entrainment air velocity (Uea sg ) needs to be greater than the terminal velocity (Utv) of the oxygen carriers. The detailed design criteria indicate the various operating parameters that need to be known prior to design and construction of the subpilot unit. One of the major design variables is the oxygen carrier properties. The following section discusses the properties of the oxygen carrier used for the subpilot scale SCL unit operation. 2.4. Oxygen Carrier for the Subpilot SCL Unit. The oxygen carrier performance and the unit design are the two key parameters that are of close relevance to the success of the SCL process. The focus of this section is on the properties and the large scale production of the oxygen carriers. The oxygen carrier moves through individual reactors undergoing continuous change because of oxygen transfer. The proper function of the oxygen carrier over multiple cycles is of importance for the economic feasibility of the SCL process. The advantage of generating H2 using steam oxidation is unique to iron-based oxygen carriers. Therefore, the oxygen carriers used for the SCL process are composite oxygen carriers with iron oxide as the active component. More than 600 different oxygen carriers have been tested for the chemical looping combustion including many iron-based oxygen carriers.45 Most of the tested oxygen carriers were designed for fluidized bed operating mode that enforces certain different requirements on the oxygen carrier.46 The SCL unit operates using a moving bed reducer and oxidizer that prevents the direct utilization of the oxygen carriers tested for other fluidized bed based CLC systems. Moreover the possibility to generate H2 calls for additional constraints on the selection of the oxygen carrier. Iron-based oxygen carriers with varying

(n)

1 O2 → FeO (o) 2 Only the first reaction is expected to be prevalent in the combustor under a H2 generation mode; while the other reactions are also important for the heat generation mode. It is clear that the objective of the combustor is to replenish the oxygen carrying capacity of the particles. The oxidizing gas used in the combustor is air. The minimum amount of air required for the complete oxidation of the oxygen carriers can be calculated based on the oxygen demand in the combustor. The oxygen demand in the combustor will vary depending on the mode of operation of the unit. In the combustion mode, the oxygen carriers exiting the reducer will be introduced to the combustor directly. Therefore, the oxygen demand will be governed by the oxygen carrier conversion exiting the reducer and the solid circulation rate. (15)

The oxygen demand in the combustor is the same as the oxygen required in the reducer to convert the fuel, which is then fully regenerated in the combustor. Under a H2 generation mode, part of the particle oxidization is performed during the steam oxidation carried out in the oxidizer. This reduces the oxygen demand in the combustor, which is given by the following equation: H2 oxi Θcom = χOC ṁ OCΨOC = Θred − Θoxi

4

Air is the reactive gas introduced into the combustor. Therefore, the air flow requirement (Qair), in SLPM, can be estimated from knowing the oxygen demand in the combustor and the excess air percentage (∈).

Fe +

red ΘΔcom = χ OC ṁ OCΨOC

(ϕcom)2

(17)

(16)

From the knowledge of the oxygen demand, the air requirement and reactor dimensions can be derived, in a 2295


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Figure 1. Schematic diagram of the SCL subpilot unit for hydrogen production on the left (a) and for combustion on the right (b). Sections 1, 3, and 4 represent the moving-bed reducer, moving-bed oxidizer, and fluidized-bed combustor, respectively. Section 2 is the rotary solids feeder. Section 5 is an entrained bed riser, and section 6 is as cyclone-based solid−gas separation device.

supports, metal oxide loadings, sintering temperatures, particle morphologies, and particle preparation techniques were synthesized and tested for moving bed based systems. More details regarding particle testing equipment and results can be found in previous studies.32,47

V (ρP − ρ)g = Cd

π μ2 2 Re tv 8 ρ

(21)

Using air as the entrainment gas, Utv is calculated to be 33.39 m/s at 900 °C. For moving beds, the reactor diameter needs to be at least ten times larger than the particle diameter for smooth solids movement. This means that the reactor diameter needs to be greater than 1.77 in. (0.045 m). 3.1. Subpilot Syngas Chemical Looping Reactor Specification. The SCL subpilot unit is designed to demonstrate both the combustion mode and the H2 generation mode. The design specification of the combustion mode is discussed first. The combustion mode involves the design of two reactors namely the reducer and the combustor. The reducer is operated as a counter current moving bed, and therefore, the gas velocity should be less than 3.67 m/s. Previous bench scale studies coupled with simulation results indicated that an oxygen carrier residence time of 30 min in the reducer was sufficient to achieve the maximum possible oxygen carrier conversion of 50% when using syngas as the fuel. Using this information, the diameter for the reducer was selected to be 3 in. (0.0762 m). This resulted in a reactor height of 10.82 ft (3.30 m) for a capacity of 25 kWth. This provides enough room for housing other equipment required to regulate solid flow as the maximum total height allowed is 22 ft (6.70 m). The superficial gas velocity calculated for the above dimensioned reducer with

3. DESIGN SPECIFICATION OF THE SUBPILOT SCL UNIT It is important to adhere to the criteria mentioned in the previous section for achieving a smoothly operating unit. Specifically, the gas flow rate used in the moving bed reducer and oxidizer should be less than the minimum fluidization velocity (Umf) to prevent axial mixing of the particles. Additionally, the gas flow rate in the combustor should be below the Umf during fixed bed operation and above the terminal velocity to convey the solids to the reducer after full oxidization. The Umf of a solid can be calculated using the Ergun equation.48 The gas properties for the reducer are based on a 1:2 molar ratio of H2/CO as the reactant gas feed while the oxidizer and combustor are based on steam and air, respectively. The operating temperature is 900 °C. The Umf is calculated to be 3.67 m/s, 3.76 m/s, and 3.10 m/s for the reducer, oxidizer, and combustor respectively. The viscosity correlation for the gas composition is calculated using Sutherland’s equation.49 The terminal velocity (Utv) of the particles is calculated by 2296


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25 kWth capacity is 1.39 m/s at 900 °C, which is significantly lower than the Umf. The solid circulation rate corresponding to 25 kWth capacity is 752 g/min under the assumption of 50% oxygen carrier conversion in the reducer. The combustor is operated in a fixed bed mode, followed by the entrainment of the oxygen carriers. A smaller reactor diameter is preferred in the riser section of the combustor, as the entrainment velocity required is 33.39 m/s at 900 °C. Therefore, the riser section of the combustor is designed with a diameter of 2 in. (0.0508 m). The reaction section diameter of the combustor was 4 in. (0.102 m). A residence time of 5 min is used for this section. Therefore, the height of the combustor to handle the solid inventory of 3.76 kg is 1.01 ft (0.308 m). The total quantity of air corresponding to the oxygen demand is 1133.8 L, which is introduced over a period of 5 min. Allowing for 10% excess air utilization, the airflow rate is 250 SLPM, which corresponds to Usg of 2.05 m/s, lower than the Usf of 3.1 m/s. The transition from reducer bottom to the reaction section of the combustor reactors and the solid flow rate is controlled by a rotary disk solid feed and two ball valve setup. The rotary solid feeder and ball valve setup are designed for a 3 in. (0.0762 m) i.d. reducer. The ball valves are more expensive as the size increases, and the solid feeder is a unique system designed specifically for the SCL subpilot unit. The riser section of the combustor is connected to the top of the reducer through a cyclone and two ball valve setup. The cyclone is designed for a transition from a 2 in. (0.0508 m) section to a 3 in. (0.0762 m) section, and 3 in. (0.0762 m) ball valves are used. The total height added to the system by the addition of the transition system is less than 10 ft (3.05 m), which allows the total height of the reactors to be slightly less than 22 ft (6.70 m). In the H2 production mode, the introduction of the oxidizer requires another transition system consisting of two ball valves and a solid feeder. The height occupied by the cyclone, the two solid feeders, and the six ball valves limit the combined height to 80 in. (2.03 m) available for the reducer and oxidizer. Therefore, the reducer is designed with a 3 in. (0.0762 m) diameter and a 40 in. (1.02 m) height. This limits the maximum solid circulation rate to 230 g/min, assuming a 30 min residence time in the reducer. Therefore, the maximum capacity of the unit under the H2 generation mode is 7.67 kWth under a 50% oxygen conversion case. The combustor is expected to handle a solid inventory of 1.15 kg when a residence time of 5 min is provided for the fixed bed combustion. Therefore, a combustor with a 2 in. (0.0508 m) diameter would result in a reaction section height of 6.6 in. (0.168 m). The air flow rate, calculated based on the oxygen demand and providing 10% excess, is 17 SLPM. Therefore, a combustor with both the reaction and riser section of 2 in. (0.0508 m) diameter can be used as the Usg during fixed bed stage is lower than the Umf. 3.2. Construction of the Subpilot SCL Unit. Based on the reactor dimensions, the schematic diagram of the subpilot SCL unit for H2 generation mode and combustion mode are shown in Figure 1. Figure 2 is a photograph of the subpilot scale SCL unit located at the Ohio State University (OSU) Energy Research Center facility. During the operation of the subpilot unit, the oxygen carrier particles travel downward from the reducer to the oxidizer reactor in a plug flow moving bed design, while the gas feed for each section travels in the opposing upward direction. Rotary feeders are placed below the oxidizer and reducer reactors controlling the solids flow rate for

Figure 2. Photograph of assembled subpilot SCL unit.

each reactor independently. A transitional section allows the particles to travel from the oxidizer to the combustor reactor. The combustor operates as a semibatch fixed and entrained bed reactor to fully reoxidize a batch of oxygen carrier particles and convey them to the reducer. The reducer and oxidizer both operate as moving bed reactors with similar residence time requirements. Therefore, the designs of these two sections are identical. The reactor consists of a 3 in. (0.0762 m) i.d. pipe with multiple ports for temperature, pressure, and gas sampling sensors of the reactor pipe. Significant heat loss is expected for this system because of the small scale of the unit. As such, external heating elements are placed along each of the main reactor sections to generate isothermal operating conditions. A set of ball valves are positioned between each of the three reactor sections to allow solids transport while mitigating gas mixing between reactor sections via a lock-hopper type operation. The ball valve transport system consists of 2−3 in. (0.0762 mm) i.d. flanged, metal-seated ball valves attached in series. Three sets of these valves are used between each reactor section to transport the solids from one section to the next while preventing gas mixing. Figure 3 illustrates the automation scheme for one set of valves. The reactant syngas feed used in the reducer is simulated using gas cylinder tanks consisting of CO, CO2, H2, CH4, and N2, as illustrated in Figure 1. A gas mixing panel designed in collaboration with Air Products is used to blend the gases to the desired composition. Separate lines of N2 and He gases are also incorporated in the gas mixing panel. A mass flow controller is installed on each gas line to regulate the gas flow. Pneumatic solenoid valves are used upstream of the mass flow controllers. Check valves are also incorporated to avoid gas back flow. The flow rate of individual gases can either be altered by a touchscreen human−machine interface (HMI) control installed on the front of the gas panel or modified remotely from a computer located inside the control room. A bypass line is assembled to the mixed outlet gas to allow gas sampling prior to being injected into the reducer. Several safety features are built into the gas panel. These features include an initial flushing sequence to eliminate the oxygen in the reactor, a 2297


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acquisition board from Measurement Computing, solid state relays, and inline fast-acting fuses. A proportional on−off control is adopted for the heating elements. This scheme is sufficient for controlling the process temperature, as the thermocouples measuring the internal solids temperature are embedded within the reactor. The heat capacity of the reactor wall and solids within prevent temperature spike readings as the heaters are turned on and off. The high frequency temperature acquisition coupled with the lower frequency heater output signal allows for effective temperature control. Energizing and de-energizing control valves in the program are based on inverted binary code. An automated loop is programmed to properly open and close each valve in a given time sequence. Programmable safeguards, such as valve locking and emergency shutdown, are also incorporated into the control system.

Figure 3. Operation scheme for updated valve assembly. Step 1: upper valve opens transferring oxygen carrier solids from upper reactor section to intermediate chamber. Step 2: upper valve closes. Step 3: lower valve opens transferring solids from intermediate chamber to lower reactor section. Step 4: lower valve closes.

4. TEST RUN AND RESULTS Two experimental studies are presented to analyze the reducer and oxidizer moving bed reactor performance. The following sections describe the experimental setup for the two test conditions with results and discussion following. 4.1. Experimental Setup. Table 1 summarizes the two test run conditions. The solids flow rate used are measured prior to

reactor temperature monitor to prevent the formation of explosive gaseous mixture, a reactor pressure monitor to prevent pressure build up in the reactor, a ventilation monitor to prevent the combustible gas flow during vent failure, and a hazardous gas monitor to prevent the discharge of hazardous gas to the environment. Multiple gas sampling ports are placed at the oxidizer and the reducer. A small slipstream of gas (less than 1 SLPM) is drawn into a micro-gas chromatograph (GC) for characterization. Prior to the analysis, gas conditioning is required to remove any fine solids and to condense steam from the sample to protect the sensitive equipment. A 2 μm sintered metal filter is placed on the end of each sampling port for fines removal. A desiccant bed is used at the end of the sampling lines for gas conditioning, while a suction pump and needle valve are used to control the gas flow rate to the analyzer. Figure 4 summarizes the gas sampling system used.

Table 1. Reducer Gas Composition solids flow (g/min) reactor operating temp. (°C) CO flow (L/min) H2 flow (L/min) calculated solids conversion (%) duration (h)

test 1

test 2

87 850 2 1 13 4

150 900 10 5 39.6 2

each experiment. The active metal loading (δ) of the particles used for demonstrations are 60 wt % iron oxide. With the syngas and solids flow rate coupled with the active particle composition and properties, eqs 3, 4, and 5 are used to calculate the theoretical solids conversion. In addition to operating criteria, a reactant and entrainment air flow of 102.66 SLPM and 2407 SLPM, respectively, were used in the combustor to fully oxidize the particles and ensure their conveyance to the reducer. Steam was injected into the oxidizer at a flow rate of 2.3−3.2 SLPM of water. The ball valve solids transfer system was automated to cycle at 2-min intervals throughout each demonstration. Two gas analyzers were used to profile the product streams of the reducer and oxidizer. A Varian CP-4900 inline microGC is used with a 3-min sampling interval as well as a Sick Maihak S71 to provide instantaneous gas composition using thermal conductivity and infrared-based analysis. For each experiment, the reactor system was brought to steady operating temperatures and safeguards checked prior to syngas injection. Continuous outlet gas profile was performed with the gas analyzer throughout each demonstration. Based on the outlet gas analysis, steady state operation was achieved prior to system shutdown. Post-experiment solids analysis was also performed after the completion of test 2. 4.2. Results and Discussion. Prior to the unit demonstration, the solids flow rate was calibrated based on the rotational speed of the solid feeder. A trend line was drawn by studying five solid rotational speeds averaged over three particle collection durations, and one repetition for each test

Figure 4. Schematic diagram of gas sampling design.

To control the ball valve assembly, external heaters, and bed height monitoring system, a process control system is developed. The software used in the previous lab-scale demonstration, DacFactory by Azeotech Inc., is adopted as the process control software. The hardware system consisted of a process control and relay board from Labjack, a temperature 2298


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caused by motor speed control loop, the effect of temperature on the solids flow rate can be considered minimal. After test 1 reached steady state operation, the gas analysis from the reducer gas outlet had an average purity of 99.88% CO2 corresponding to 99.96% syngas conversion. In addition to the high syngas conversion, harmonious system automation of the subpilot unit under high temperatures was accomplished. The calculated reducer solids conversion for test 1 is 13%, slightly greater than the thermodynamic limit for H 2 production from the oxidizer. Steam was injected into the oxidizer during the 4-h demonstration, and the gas profile of the oxidizer was analyzed using the GC. The results indicate that hydrogen production was restricted as only small amount of H2 could be seen from the gas analysis. The low solids conversion from the reducer mitigates the capability of the oxidizer in producing H2 from the steam−iron reaction. Single stage fluidized-bed chemical looping combustion with iron oxide processes have a similar disadvantage as the solids conversion in the reducer must be maintained below 11.11% in order for full syngas conversion. Previous 2.5 kWth bench scale studies demonstrated the capability of the moving bed system to generate a solids conversion of nearly 50% while maintaining high syngas conversion.14,19 Such particle reduction allows for favorable hydrogen production from the oxidizer. For test 2, the calculated solids conversion is 39.6%. The system solids flux was increased from the previous to provide a higher hydrogen production from the oxidizer. After the system was brought to the operating temperature, syngas and steam were injected to the reducer and oxidizer, respectively. The average syngas conversion during the first 40 min of test 2 was 99.5% with the highest conversion being 99.99%. The syngas conversions in the reducer obtained from the microGC are consistent with the results from the Sick Maihak gas analyzer. Figure 8 illustrates the reducer gas profile for test 2. After a steady gas composition was achieved in the reducer operation, gas sampling was switched to oxidizer gas outlet. Figure 9 illustrates a 30-min snapshot of the oxidizer gas profile. The average hydrogen purity generated from the oxidizer is 94.4% with a maximum of 98.4%. After the subpilot unit was shutdown and brought to room temperature, a solids sample was taken from below the oxidizer and reducer. The samples were analyzed to determine the solids conversion from each reactor section. The conversion was found using a thermogravimetric analyzer (TGA) to determine the weight change in the sample associated with fully reoxidizing the solids to Fe2O3. The solids discharged from the reducer were found to have a solid conversion of 35.54% resulting in iron particles entering the oxidizer in the FeO/Fe oxidation state. The calculated theoretical value of the solids conversion is consistent with the experimentally determined value. The solids conversion for the sample discharged from the oxidizer was found to be 22.36%. The experimental study of the subpilot SCL unit successfully demonstrated the integrated 3-reactor performance with the potential for the cogeneration of purified streams of CO2 and H2 from gasified coal. The moving bed reactor scheme is capable of high syngas conversions while producing sufficient solids reduction in the reducer allowing for hydrogen generation from the oxidizer.

run. Figures 5 and 6 illustrate the calibration results for the upper feeder, controlling the reducer solid flow rate, and the

Figure 5. Relationship between motor rotational speed to solids mass flow rate for the top solid feeder controlling the reducer at room temperature.

Figure 6. Relationship between motor rotational speed to solids mass flow rate for the bottom solid feeder controlling the oxidizer at room temperature.

lower feeder, controlling the oxidizer solid flow rate, respectively, at room temperature. Figure 7 illustrates the

Figure 7. Relationship between motor rotational speed to solids mass flow rate for the top solid feeder controlling the reducer at reaction temperature.

results for the upper solids feeder calibration when operating the reducer at the expected reaction temperature. Comparing Figures 7 and 5, the resulting trend has a 4% difference in solids flow rate to disk rotational speed at a 900 °C operating temperature when compared to the calibration at ambient conditions. Taking into account solid flow rate fluctuations 2299


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Figure 8. Syngas conversion during the first 40 min of reducer operation in test 2.

Figure 9. H2 product purity from the oxidizer during test.

5. CONCLUSIONS

continuously generating 94.5% hydrogen with in situ carbon capture using the SCL process. The higher oxygen carrier conversion (35.54%) attained as a result of using a moving bed reducer is associated with the ability to generate hydrogen. Very high syngas conversion (>99%) is observed in all the tests to confirm the ability to achieve in situ carbon capture. The results support the possibility of utilizing SCL process for high purity hydrogen generation from coal with in situ carbon capture.

The SCL process presents the unique opportunity to cogenerate hydrogen and electricity with in situ carbon capture by the cyclic reduction and regeneration of iron-based oxygen carriers using syngas and steam/air, respectively. The novel reactor configuration allows for the product flexibility associated with the SCL process. The design criteria for both the moving bed reactors namely reducer and oxidizer have been discussed in detail. The influence of the oxygen carrier properties on the reactor design is elaborated. The 25 kWth subpilot unit was designed and constructed based on the information available from prior results and the design criteria. The initial operational results highlight the possibility of

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AUTHOR INFORMATION

Corresponding Author

*Telephone: (614)-688-3262. Fax: (614)-292-3769. Email:fan. [email protected]. 2300


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Present Address

OC = oxygen carrier

†

Currently with North Carolina State University

Superscript

Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors would like to acknowledge fellow graduate student, Fei Wang, for his very helpful support in the reactor design and demonstration. The authors would also like to acknowledge the financial assistance provided by the Ohio Coal Development Office (OCDO) of the Ohio Air Quality Development Authority (OAQDA) and Third Frontier Advanced Energy Program (TFAEP) of the Ohio Department of Development (ODOD) in support of the SCL subpilot unit construction and demonstration are gratefully acknowledged.

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r = reducer solid outlet o = oxidizer solid outlet ea = entrainment air com = combustor red = reducer oxi = oxidizer Δ = combustion mode H2 = hydrogen generation mode

REFERENCES

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NOMENCLATURE

Notation

m = mass (g) g = acceleration due to gravity (m/sec2) Q = volumetric flow rate (SLPM) C = capacity of the unit (kWth) V = volume (m3) H = height (m) P = pressure (Pa) U = velocity (m/s) R = oxygen mass ratio T = temperature (K) η = efficiency (%) ϕ = diameter (m) Θ = oxygen demand (g/min) χ = conversion (%) Ψ = oxygen carrying capacity δ = active metal oxide loading ṁ = mass flow rate (g/min) τ = residence time (min) ρ = density (kg/m3) ∈ = excess air ratio μ = viscosity (kg/m.sec) Cd = drag coefficient Re = reynolds number Subscript

f = final g = gasifier b = bulk P = particle R = reactor r = maximum reduced weight ox = maximum oxidized weight th = thermal tv = terminal velocity sg = superficial gas velocity mf = minimum fluidization velocity syn = syngas red = reducer oxi = oxidizer com = combustor air = air flow max = maximum H2O = steam CO = carbon monoxide H2 = hydrogen 2301


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