Wang Du, Liping Ma,*, Jing Yang, Wei Zhang, Ran Ao

1 Faculty of Environmental Science and Engineering, Kunming University of Science and Technology, Kunming 650500, China

2 School of Environmental and Municipal Engineering, Xi’an University of Architecture and Technology, Xi’an 710000, China

Keywords:Phosphogypsum Lignite Chemical looping gasification Fluidized-bed Syngas Computational fluid dynamics

ABSTRACT Phosphogypsum(PG)is a solid waste produced in the wet process of producing phosphoric acid.Lignite is a kind of promising chemical raw material.However,the high sulfur of lignite limits the utilization of lignite as a resource.Based on fluidized bed experiments,the optimal reaction conditions for the production syngas by lignite chemical looping gasification (CLG) with PG as oxygen carrier were studied.The study found that the optimal reaction temperature should not exceed 1123 K;the mole ratio of water vapor to lignite should be about 0.2;the mole ratio of PG oxygen carrier to lignite should be about 0.6.Meanwhile,commercial software Comsol was used to establish a fuel reaction kinetics model.Through computational fluid dynamics (CFD) numerical simulation, the process of reaction in fluidized bed were well captured.The model was based on a two-fluid model and coupled mass transfer, heat transfer and chemical reactions.This study showed that the fluidized bed presents a flow structure in which gas and solid coexist.There was a high temperature zone in the middle and lower parts of the fluidized bed.It could be seen from the results of the flow field simulated that the fluidized bed was beneficial to the progress of the gasification reaction.

1.Introduction

Syngas is mainly composed of H2and CO produced by natural gas or solid fuels,and is an important raw material for the chemical energy industry.Especially it plays an important role in the chemical industry such as synthetic ammonia and methanol.Nowadays,the steam reforming of natural gas is the main method in the production of syngas.However, it requires expensive air separation equipment.In recent years, it has been proposed to use chemical looping gasification (CLG) technology to produce syngas [1,2].

CLG is a new type of gasification technology, which developed from chemical looping combustion (CLC) [3].CLG is characterized by not directly using oxygen molecules in the air,but using lattice oxygen in the oxide to complete the fuel combustion process,which avoids direct contact between the fuel and the air [4,5].The principle is shown in Fig.1.In the fuel reactor,the fuel is gasified, the oxygen carrier and the fuel undergo an oxidation-reduction reaction to produce synthesis gas, and the oxygen carrier loses oxygen to produce MexOy-1; In the air reactor, MexOy-1reacts with air to regenerate the oxygen carrier[6].Compared with traditional coal gasification, CLG has many advantages.First of all,CLG technology avoids expensive air separation equipment and saves costs.Secondly, the oxygen carrier can be recycled and regenerated, and the gasification reaction can be continued without external heating.

Phosphogypsum (PG) is a hazardous solid waste and a byproduct of the wet production of phosphoric acid, and its main component is CaSO4[7,8].CaSO4has been proven to be a promising oxygen carrier[9-12].According to our team’s research,PG can be directly used as an oxygen carrier [2,13].PG is used as an oxygen carrier, which not only effectively alleviates the accumulation of PG, but also benefits the environment.Lignite is the coal with the lowest degree of coalification.It has the characteristics of high moisture, high volatility, low calorific value and high sulfur content.Lignite can be used as a solid fuel [14], which is extremely important for building an environmentally friendly society.

Fig.1.Schematic illustration of the chemical-looping gasification process.

The reactor plays an important role in the process of the chemical looping reaction.The research of chemical looping reactor has gone through two stages of fixed bed and fluidized bed.With the development of fluidization technology, gas-solid fluidized beds have been widely used in the chemical,petroleum and other industries.During fluidization, the gas-solid phase can be fully contacted, and the relative movement between the particles and the fluid is frequent and intense, which makes the heat and mass transfer efficient.Lyngfelt used a serial fluidized bed reactor to carry out a pilot test of chemical looping combustion, and the results showed that the fuel conversion rate was as high as 99.5%[15].In Wheelock’s double atmosphere fluidized bed experiment,the decomposition rate of PG could reach 97% [16].The fluidized bed system built by Swift had a solid desulfurization rate of 98%due to its recyclable characteristics[17].According to many documents, the fluidized bed system could significantly improve the physical and chemical properties of solid particles [18-20].Therefore, fluidization technology has great advantages in gas-solid reactions.

In view of the complexity of the gas-solid fluidized bed twophase flow structure, and the flow mechanism is not very clear,so the research method has been based on experimental means for many years.With the improvement of numerical simulation calculation methods, the time required for Computational Fluid Dynamics (CFD) is decreasing [21].Models for complex gas-solid multiphase flow, mass transfer, heat transfer, and chemical reaction become more accessible.Generally, in the CFD simulation of gas-solid flow, Lagrangian and Eulerian approaches are widely used.In the Eulerian approach, such as the two-fluid model(TFM), both the gas phase and the particle phase are considered continuous phases, and the particles and fluid are coexistent and interpenetrating continuous media.Currently the TFM based on particle dynamics theory has been widely used in numerical simulation of heat transfer in fluidized bed [22-25].In the TFM, each phase is described by separate conservation equations, and interphase interactions are characterized by interphase coupling.The solid phase is treated as a single phase, resulting in a significant reduction in the number of equations to be solved, which saves computational costs.In recent years, numerical simulation has become a promising tool to simulate gas-solid two-phase flow.The calculation of different models is used to study the performance of the fluidized bed system [26-31].However, in most of the previous fluidized bed numerical simulation studies, there are not many literatures using CFD simulation to calculate the chemical looping gasification process.

In this paper, the factors affecting the formation of syngas by lignite CLG with PG as oxygen carrier were studied through fluidized bed experiments.The effects of various factors on syngas were explored.Meanwhile,the fuel kinetic model was established based on Eulerian approach to simulate the behavior of reaction flow in reactor.The characteristics of CLG in a fluidized bed were discussed by CFD simulation method.

2.Experimental and Numerical Simulation

2.1.Experimental material

The oxygen carrier used in the experiment was PG,and its composition are shown in Table 1.The reducing agent used was lignite,and its composition is shown in Table 2.The oxygen carrier sample was dried in a 378 K air drying box for 2 hours, and then passed through a 80-mesh sieve.The lignite sample was crushed and screened through a 120-mesh sample sieve.According to the requirements, different proportions of PG and lignite mixtures were prepared.The particle size of the sample was analyzed by Mastersizer 3000 laser particle size analyzer.The average particle size of the mixture is 48 μm, and average particle size of oxygen carrier and lignite is shown in Fig.2.

Table 1 Composition of phosphogypsum sample (%, mass)

Table 2 Composition of lignite sample

2.2.Experiment procedure

The experiment was carried out in a fluidized bed reactor.The device is shown in Fig.3,which is mainly composed of steam generator, air supply system, feeding system, riser, gas-solid separation system, return box, etc.The reactor has a shape with a wide top and a narrow bottom.During the fluidized bed reactor was brought to the specified temperature, the fluidized bed reactor was heated by resistance wire.The nitrogen was provided by gas cylinder and the water vapor was provided by the steam generator.Water vapor was used as a fluidizing air.One hour after the introduction of nitrogen, a single feed method was adopted, that is,before the reaction began, the fluidized bed reactor was raised to a specified temperature, and then mixed oxygen carriers and lignite were added at one time.Then the steam generator was turned on.The water vapor flow rate was controlled by the amount of water entering the steam generator.During the reaction, the fluidizing gas entered the lower port of the reactor, and the upper end was provided with an outlet with a diameter of 60 mm.After the reaction, the light component flowed out from the upper port,and the heavy component flowed out from the lower port.The mixture was purged with a certain proportion of nitrogen and steam to make the mixture quickly fall into the reactor.There was a rapid and complex reaction between oxygen carrier and fuel lignite.The generated gas passed through a silica gel drying bottle and then passed into the flue gas analyzer for detection.

The calculation formula of carbon conversion rate (Xc) is as follows.

where VCO2, VCO, VCH4and mcare the volumes of CO2, CO, CH4and the carbon content in lignite, respectively.

2.3.Numerical simulation

2.3.1.Mathematical model

The CFD numerical simulation of the fluidized bed was carried out.The main assumptions made by the model were:

(1) Both the gas phase and the particle phase were regarded as continuous phases in the whole process.It was considered that particles and fluid are coexistent and interpenetrating continuous media.

(2) The particles in the fluidized bed were spherical and uniform in size.

(3) The general reaction equation of the oxygen carrier and brown coal gasification process was simplified as:

Fig.2.Average particle size of oxygen carrier and lignite.

Fig.3.Schematic diagram of fluidized bed (mm).

2.3.2.Equations

(1) Continuity equations

where u,v and w are the speeds in the x,y,and z directions,respectively,and ρ is the fluid density.The mass conservation equation is applicable to any fluid problem.Its law can be expressed as: The increase in the mass of the fluid micro-element body per unit time,that is, the net mass flowing into the micro-element body in the same time interval, from which the conservation equation is derived.Introducing vector symbols div（a）＝,Eq.(3) is transformed into:

(2) Momentum equations

where p is the pressure on the fluid micro-element;τxx,τxy,and τxzare the components of the viscous stress τ; Fx, Fy, and Fzare the physical forces on the micro-element body.

(3) Energy equations

where cp,T,k,and sTare the terms of specific heat capacity,temperature, heat transfer coefficient, and viscosity, respectively.The law of conservation of energy is a basic law that must be satisfied by a flow system that includes heat exchange.

(4) Species transport equations

where cs, Ds, Ssare the volume concentration, diffusion coefficient,and the amount of chemical reaction s component produced per unit time and unit volume, respectively.In a specific system, there are multiple chemical components,and each component must comply with the conservation equation of components.It can be described as:the rate of change of the quality of a certain chemical component in the system with respect to time is equal to the sum of the net diffusion flow through the system interface and the production rate of the component produced by the chemical reaction.

2.3.3.Simulation method

The two-fluid model was established by Euler-Euler method.The conservation equations of fluid phase and particle phase are established in Euler coordinate system.The differential equations were all solved by finite volume method.The upwind difference scheme was used to carry out the discrete equation on the finite volume,and Comsol was used to solve it.The convection term useda first-order discrete format.The phase coupling adopted the SIMPLE algorithm.The specific parameters of the simulation process are shown in Table 3.

Table 3 Simulation model parameters

The RNG k-ε model proposed by Yakhot and Orzag [32]was selected.This model can be used to better handle flows with high strain rates and large streamline curvatures.The coefficient of restitution between solid particles was 0.9.The heat transfer coefficient between the gas phase and the solid phase proposed by Gunn [33]was selected.The shrinking-core model was adopted,which has been successfully verified in the study of particle reactivity [34].For the boundary conditions, velocity inlet, pressure outlet, and constant temperature adiabatic wall were selected.No sliding wall condition was used for gas and solid phases.The Syamlal-O’Brien drag force model [35]was selected as the gassolid drag force.The bed was initially filled with solid particles with a height of 0.2 m, and the total volume fraction was 0.63.There were 2000,000 hexahedral discrete units in the 3D computational domain, as shown in Fig.4.

Fig.4.The gird of simulation calculation.

Fig.5.Effect of reaction temperature on gas phase products.

3.Results and Discussion

3.1.Experimental results and discussion

3.1.1.Effect of temperature on gas production

Fig.6.Effect of reaction temperature on carbon conversion.

Fig.5(a) shows that as the temperature increases, the time required for the CO gas reaction was significantly reduced.When the reaction temperature was 1023 K, CO was produced completely within 30 minutes.When the temperature rose to 1073 K, CO production was mainly concentrated in the first 20 minutes.When the temperature continued to increase to 1123 K,CO production was mainly concentrated in the first 12 minutes.This showed that when the reaction temperature was in the range of 1023-1123 K, temperature was the main controlling factor affecting CO production.The reason was that increasing the temperature was beneficial to the coal gasification reaction.In addition, the increase in temperature was conducive to activate the active sites on the carbon surface, thereby increasing the carbon conversion efficiency [36,37].Compared with 1123 K, when the reaction temperature increased from 1123 K to 1173 K, the maximum CO production rate was reached faster at 1127 K.

Fig.5(b) shows the distribution of CO2gas at different temperatures.CO2production was mainly concentrated in the 5th to 20th minutes of the beginning of the reaction.And with the increase of temperature, CO2production gradually increased.This showed that high temperature was favorable for the oxygen carrier to oxidize most of the gasification products to carbon dioxide.

Fig.5(c)shows the amount of H2produced at different reaction temperatures.The release rate of H2reaches its peak within 5 minutes.At different reaction temperatures, the higher the reaction temperature, the lower the rate of H2generation.This conclusion was verified by relevant literature[38,39].This was due to the fact that higher temperature were conducive to the occurrence of water gas reaction,when the reaction temperature was too high,the oxygen carrier oxidation reaction of most gasification products was more dominant than the water gas reaction [40].When the reaction temperature increased from 1023 K to 1123 K, the amount of CO production increased steadily.However, when the temperature rose from 1123 K to 1173 K,the increase of CO production was very slow.At the same time,as the reaction temperature increased from 1023 K to 1173 K, the amount of H2produced gradually decreased, especially at 1123-1173 K, the production rate of H2was very low.With the continuation of the reaction, the reduced oxygen carrier was wrapped around the outside of the oxygen carrier, which would block the oxygen carrier from contacting the vaporized product.The contact time and contact surface of the gasification product and the oxygen carrier decreased, causing the H2release rate to gradually decrease to zero.The results showed that the reaction temperature of the fluidized bed test should not exceed 1123 K.

Fig.6 shows the conversion of carbon at different reaction temperatures.As the reaction temperature gradually increased from 1023 K to 1073 K, the carbon conversion rate increased slowly when it exceeded 1123 K.This was due to the higher reaction temperature was beneficial to the water gas reaction.However, when the temperature exceeded 1123 K, the excessively high temperature was more conducive to the process of oxygen carrier reduction by gasification products [41].As a result, the rate of carbon consumption will decrease.

Fig.7.Effect of different molar ratio of phosphogypsum to lignite values on gas phase products.

3.1.2.Effect of oxygen molar ratio of PG to lignite on gas production

Fig.7 shows the changes in gas phase products with four different molar ratio of PG to lignite.With the increase of the oxygen carrier ratio,the production of CO and H2increased steadily.When molar ratio of PG to lignite was greater than 0.7,the production of CO and H2begins to decrease.This was because in the reaction process,the iron-based oxide contained in the PG oxygen carrier were first reduced by the gasification product.The reduced iron compounds catalytically reduced CaSO4to CaS, thereby promoting the production of CO2.However,when the molar ratio of PG to lignite was greater than 0.7, the oxygen carrier reduction reaction was more dominant than the water gas reaction, which led to the decrease of CO and H2production.

Fig.8.Effect of different molar ratio of phosphogypsum to lignite values on carbon conversion.

Fig.8 shows the carbon conversion rate at different oxygen mole ratios between PG and lignite.With the increase of the oxygen molar ratio of PG to lignite, the carbon conversion rate increased steadily and eventually stabilizes.Combined with the analysis of the influence of the oxygen mole ratio of PG and lignite on the amount of gas phase products, it was found that when the oxygen mole ratio of PG to lignite is greater than 0.7,the reduction process of oxygen carrier by gasification products to oxygen carrier was more dominant than that of water gas reaction.This will lead to a decrease in the final carbon conversion rate.Therefore, the optimal PG to lignite oxygen mole ratio was about 0.6.However,in the actual industrialized operation process, it is necessary to screen according to actual needs [42].

3.1.3.Effect of water vapor content on gas production

Fig.9 shows the instantaneous distribution of gas phase products.As the mole ratio of water vapor to lignite increased, syngas production was steadily increasing.This was due to water vapor promoted the occurrence of water gas reaction, resulting in an increase in the production of CO and H2in the system.When the ratio of water vapor to lignite material exceeded 0.3, the amount of CO production gradually decreased steadily.This reason was that the water vapor on the surface of the reactant particles was close to saturation,and it was believed that the influence of external diffusion had been eliminated.Excessive water vapor content may also increase the energy consumption of the reaction system.

Fig.10 shows the effect of water vapor content on carbon conversion.As the water vapor content increased, the carbon conversion rate gradually increased.However,when the amount of water vapor exceeded 0.3, the carbon conversion rate began to decrease which is consistent with the conclusion obtained in Fig.8.It can be concluded that the optimal mole ratio of water gas to lignite should be 0.2-0.3.

3.2.Simulation results and discussion

Fig.9.Effect of water vapor content on gas phase products.

Fig.11 shows the comparison between the simulated and experimental values of the concentration of outlet gas components(CO, H2, CO2).Compared with the experimental data, the volume fraction of CO and CO2in the simulated value was larger, and the volume fraction of H2was smaller.The reason was that, on the one hand, the empirical formula of each model in the simulation calculation may be different from the actual ratio and composition;on the other hand, certain errors were caused by the operation,control and equipment during the experiment.Due to these reasons, the experimental results were somewhat different from the simulated values.Through the comparison of numerical simulation results and experimental results, the average relative error was less than 15%, and the comprehensive model established was verified.

Fig.10.Effect of water vapor content on carbon conversion.

Fig.11.Simulation and experimental results of gas concentration at the outlet of fluidized bed.

Fig.12 shows velocity vector field,velocity contour of side view and velocity vector field of 3D view.The results showed that the velocity distribution in the reactor is quite different.Generally speaking, the lower part of the reactor had a higher speed and the upper part had a lower speed, which was mainly related to the shape of the reactor.The gas was fed in uniformly from the bottom,and the particles moved upward in the central area of the fluidized bed due to the driving action of the gas flow.However, the solid particles moved downward due to the upper particle material and their own gravity,which caused the central part of the reactor to increase with high speed.The velocity of the gas phase gradually decreased, and small eddies appeared in some places, and the velocity and flow field distribution of solid particles were similar to the gas phase distribution.The velocity distribution of the gas and solid phases in the fluidized bed conformed to the gas-solid fluidization characteristics in the fluidized bed, achieving a good fluidization effect.When the materials were in the fluidized state,they could be fully mixed,which improved the reaction efficiency.

Fig.12.Velocity vector field, velocity contour of side view and velocity vector field of 3D view.

Fig.13.Velocity vector field and velocity contour of reactor.(a) Velocity contour; (b) Velocity vector field.

Fig.13 shows velocity vector field and velocity contour of reactor at different heights.It could be found that the velocity distribution at different height sections varies greatly.The higher speed was mainly concentrated in the lower part of the reactor.In addition, due to the fluidization of the material in the reactor and the friction between the material and the side wall, the flow velocity of the material changed greatly in the same section.

Fig.14 shows the temperature distribution in the fluidized bed reactor.The fluidized bed showed a flow structure in which gas and solid phases coexist.The reaction mainly occurred in the lower middle end of the fluidized bed reactor.This was due to the heavier particles of the mixture of oxygen carrier and lignite at the beginning of the reaction, mainly concentrated in the middle and lower sections of the fluidized bed reactor.Meanwhile the gasification reaction was an endothermic reaction, and the solid particles themselves have a high reaction temperature.After the endothermic reaction occurred in the lower part of the fluidized bed, the heat would be instantly replenished.Therefore, there was a high temperature zone.

Fig.14.The temperature distribution in fluidized bed.

Fig.15.The pressure distribution in fluidized bed.

Fig.15 shows the pressure distribution in fluidized bed.The lower part of the fluidized bed and the upper outlet position in the fluidized bed presented a negative pressure state.The negative pressure in the lower part was because the reactant particles were heavy components,mainly concentrated in the lower half.The particles moved upward due to the driving effect of the airflow,while the solid particles continued to squeeze downward because of their own gravity and viscous force, resulting in negative pressure in the lower half.The gas phase was mainly concentrated in the middle and upper part of the fluidized bed during the reaction.Under the action of the airflow, the gas was continuously discharged from the upper outlet,thus presenting a positive pressure state.In addition, the negative pressure at the outlet led to the reflux of the gas, which was not conducive to the timely outflow of the gas.One of the most important reasons was that when the gas reached near the outlet, some of the gas was blocked by the upper wall.The pressure on the upper wall increased,and the pressure at the middle outlet became smaller and showed negative pressure.In order to solve this problem, the cross-sectional area of the outlet should be increased.

Fig.16 shows the distribution of CO, CO2, and H2O distribution with the height of the fluidized bed, respectively.It can be seen from the figure that the generated gas was mainly concentrated on the upper part of the fluidized bed.The gas produced was a light component with a lower density than solid particles.The gas gradually moved to the upper part of the fluidized bed with the driving action of the gas flow, and was finally discharged from the upper outlet.Under the action of air flow, some of the solid particles in the middle and lower part of the fluidized bed appeared in the upper wall of the fluidized bed reactor.In addition,at high temperatures,the source of CO was mainly related to the incomplete combustion of carbon in lignite.It mainly occurred in the dense particle phase area in the middle of the furnace.The lignite particles continued to fluidize with the gas flow.The generated CO gas continued to move upwards, and with the simultaneous progress of multi-step reactions, the volume concentration of CO increased with the increase of fluidization height.

Fig.17 shows solid volume fraction distribution in the fluidized bed.The solid particles were mainly distributed in the upper middle part of the fluidized bed,and the concentration in the middle of the furnace in the fluidized bed was significantly higher than that of the side wall.Because the solid particles were relatively heavy,when the air flow was driven from bottom to top, it was mainly concentrated in the middle.At the same time, some unreacted PG and lignite particles would adhere to the side walls, making most of the solid particles concentrated in the middle and lower positions.

Fig.17 also shows that some solid particles were deposited on the upper part of the fluidized bed.This was because the slope of the ridge connecting the thin tube and the thick tube was too small, which prevented the particles from falling in time.This was not conducive to the reflux reaction of solid particles.At the same time, some solid particles were blown out at the outlet,which was not conducive to the recovery of solids from the bottom.According to the above problem, the slope of the ridge connecting the thin tube and the thick tube should be increased, and the height of the fluidized bed reaction zone should be increased.

4.Conclusions

This work is carried out from two aspects of fluidized bed experiment and numerical simulation.The characteristics and influencing factors of syngas produced by lignite chemical looping gasification with PG as oxygen carrier are examined.The following conclusions are drawn.

Fig.16.Gas volume fraction distribution.(a) CO volume fraction distribution.(b) CO2 volume fraction distribution.(c) H2 volume fraction distribution.

Fig.17.Solid volume fraction distribution at steady state.

(1) The fluidized bed experiment analysis results show that:the temperature should not exceed 1123 K; 1073 K is recommended; the molar ratio of water vapor to lignite is about 0.2; the molar ratio of PG to lignite is 0.6.

(2) The numerical simulation analysis results show that: the oxygen carrier and lignite gasify well in the fluidized bed,and the fluidized bed shows a flow structure in which solid phase and gas phase coexist.There was an obvious high temperature zone in the middle and lower part,and this temperature distribution was extremely conducive to the progress of the reaction.After the flow field was stable, the overall distribution of the multiphase components was: solid particles are mainly distributed in the middle and lower position in the fluidized bed, and gas phase components are mainly distributed in the upper position in the fluidized bed.Meanwhile, the optimization of the fluidized bed structure was that the cross-sectional area of the outlet should be increased, the slope of the ridge connecting the thin tube and the thick tube should be increased, and the height of the fluidized bed reaction zone should be increased.

(3) The simulation and experimental results of the gas concentration at the outlet of the fluidized bed show that:the comprehensive model was verified, and the calculated results were accurate, which could be used to predict the chemical looping gasification process in the fluidized bed.

Declaration of Competing Interest

The authors declare that they have no known competing financial interests or personal relationships that could have appeared to influence the work reported in this paper.

Acknowledgements

Financial support for this project were provided by National Natural Science Foundation of China (No.21666016), National Key Research and Development Program of China(2018YFC1900200) and State Key Laboratory of High-efficiency Utilization of Coal and Green Chemical Engineering (2021-K39),which is greatly acknowledged.