Renren Zhang,Yang Huang,Kaitian Zheng,Chunjian Xu*

School of Chemical Engineering and Technology,State Key Laboratory of Chemical Engineering,Chemical Engineering Research Center,Tianjin University,Tianjin 300072,China

Keywords:Fischer-Tropsch synthesis Mixed alcohols Azeotrope Separation Simulation Control

ABSTRACT Since the minimum-boiling azeotropes of C2-C8 alcohols with water and high-water content(up to 95%(mass))in the Fischer-Tropsch aqueous by-products,the separation is energy-intensive and challenging.The energy-saving strategy for the complete separation of the Fischer-Tropsch aqueous by-products has received massive attention in recent decades.In this study,a stripper-sidestream decanter process is proposed by exploiting homogeneous azeotropes (C2-C3 alcohols-water) and heterogeneous azeotropes(C4-C8 alcohols-water).The introduction of the stripping column for pre-dehydration avoids the revaporization of the mixture,and energy carried by the overhead vapor is conserved instead of being removed in a condenser.The precise fraction cutting of C1-C3 alcohol-water mixture,C4-C8 alcohols,and water is realized by the sidestream distillation column.The C4-C8 alcohols rich mixture withdrawn from the sidestream flows into the decanter to break the distillation boundary,where the organic phase returns to the sidestream distillation column to obtain the dehydrated C4-C8 alcohols,and the aqueous phase enters the stripping column.Steady-state optimization based on total annual cost(TAC)minimization shows that the stripper-sidestream decanter process reduces TAC by 17.00% and saves energy by 21.27% compared with the conventional three-column distillation process.Further,a control structure of the process is established,and dynamic simulations show that the control structure combining a differential controller with a low-selector exhibits robust control.This study provides a novel design scheme and deepens the insights into the efficient separation of aqueous by-products of the Fischer-Tropsch synthesis.

1.Introduction

The high-value conversion and clean coal utilization are strategically significant to achieving sustainable national development[1,2].Indirect coal liquefaction (ICL) technology is gradually becoming one of the important ways to decrease reliance on imported fuels and achieve efficient conversion of coal resources due to its mild operating conditions and suitability for largescale industrial production [3-5].In the production process,coal is first gasified to produce synthesis gas (the main components are carbon monoxide and hydrogen).Fischer-Tropsch synthesis converts syngas into liquid industrial products.In this work,the products from cobalt-based catalyzed Fischer-Tropsch synthesis are studied as a representative example.It includes organic products (gasoline,diesel,kerosene,and wax) and yields equivalent aqueous by-products consisting mainly of 95% (mass) water and 5% (mass) alcohols (methanol,ethanol,propanol,butanol,pentanol,hexanol,heptanol,and octanol) [6].

The separation of Fischer-Tropsch products is currently focused on separating high-value organic products,while the purification and separation of aqueous by-products are less studied [7].This complete separation of the alcohol-water mixture can recover valuable alcohol products,improve the economic efficiency of coal liquefaction,help clean utilization of coal,and contribute to environmental protection [8-13].Meanwhile,the study of aqueous by-products separation can also provide a design reference for the separation of multi-component azeotropes.

In the conventional separation process (Fig.1),the C1-C3 alcohol-water mixture,pure water,and dehydrated C4-C8 alcohols could be obtained by the fraction cutting process,thus achieving the separation of C1-C3 alcohols in the subsequent separation step.Currently,for the separation of aqueous mixed alcohols,research mainly focuses on alcohol-water azeotropic separation[14-17] rather than the fraction cutting process accounted for much higher energy consumption.

For the separation of high-water content systems,preconcentration is routinely employed to remove the water before subsequent separation.In a conventional three-column process,the dehydrating distillation and alcohol separation column accounts for 97.0%of the energy consumption of the entire process.The dehydrating distillation column is characterized by small reflux ratios,low top product purity requirements,and high wastewater concentrations at the column bottom.Based on the above,a conventional pre-concentration column could be replaced by a stripper column for energy preservation.Changet al.[18]achieved dual energy and economic saving by applying a stripping column that replaced a pre-concentrator and an azeotrope recovery column in a conventional three-column separation for aqueous isopropanol systems.Despite the use of the stripping column reduces the separation capacity and increases the processing capacity of the subsequent columns,the high retention of the vapor heat could reduce the energy consumption in the entire process.After pre-concentrationviaa stripping column,the flow has larger water content and becomes more challenging to be separated compared with using a conventional column.A larger reflux ratio and a higher theoretical stage number are required in the subsequent separation when applying the conventional strategy.However,due to the heterogeneous nature of C4-C8 alcohols,a sidestream can remove huge amounts of water without additional energy.For separating multi-component aqueous mixtures,Wuet al.[19] and Ferreiraet al.[20] proposed a strategy to obtain the target productviathe sidestream flow into the decanter while the aqueous phase of the decanter was discharged as wastewater.However,the wastewater still contains a certain amount of the target product,causing waste and contamination.Therefore,the aqueous phase of the decanter can be used as a feed to the stripping column to recover some of the mixed alcohols for the separation of the Fischer-Tropsch aqueous by-products [21].

Fig.1.Conventional separation process of aqueous by-products from the Fischer-Tropsch process.

Fig.2.(a) Ternary phase diagram and residue curve maps for water-ethanol-butanol system at 101.3 kPa.(b) The stripper-sidestream decanter process.

Fig.3.(a) Single column process for parameter optimization of the sidestream distillation column.(b) Liquid composition profiles in the single column.

Fig.4.Sequential iterative optimization sequence.

In this work,a novel separation sequence is proposed for the system of mixed alcohols from aqueous by-products of Fischer-Tropsch synthesis.The conceptual design is used in early design to verify its separation feasibility[22-26].The process is optimized based on total annual cost (TAC) for steady-state design.Further,dynamic control structures are developed for the optimized separation sequence,and the dynamic control performance between different control structures is investigated.

2.Steady-state Design of the Overall Process

2.1.Thermodynamic model and optimization objectives

The NRTL (non-random two liquids) model is used to calculate the vapor-liquid equilibrium of the alcohol-water system,and the UNIQUAC (universal quasi-chemical) model is employed to calculate the liquid-liquid phase equilibrium of C4-C8 alcohols and water [27,28].The simulated vapor-liquid equilibrium data[29,30]and liquid-liquid equilibrium data[31]of the alcohol-water system basically agree with the queried experimental data[32]which proves that the simulated data are highly credible.

The feed flow rate is 22000 kg·h-1.Nine components are taken into account to simulate a realistic composition of aqueous byproducts from the Fischer-Tropsch process.The mixture contains 94.38% (mass) water and 5.62% (mass) C1-C8 alcohols,and the feed temperature is 298.15 K.It is specified that the butanol concentration of the C1-C3 alcohols-water mixture product is not more than 0.01%(mass),the alcohol concentration of the extracted pure water is less than 0.0001% (mass),and the water content of the C4-C8 alcohols is less than 0.01% (mass).The TAC minimization proposed by Douglas [33] is applied as the objective function for screening the optimal design parameters.The TAC includes operating costs and equipment investment costs,and its calculation can be simplified as follows:

Fig.5.(a) Effect of NS on reboiler duty of the sidestream distillation column.(b) Effects of NT1 and NT2 on TAC.

Fig.6.Flowsheet of the stripper-sidestream decanter sequence.

Fig.7.Flowsheet of conventional three-column sequence.

Herein,OC is the operating cost,mainly including steam cost,cooling water cost,and electricity cost;FCI represents the fixed capital investment,including column shell,column stage,and heat exchanger construction cost;andiris the fixed capital recovery rate applied to FCI (assumed to be 0.3 here).

2.2.Conceptual design of a ternary system: water-ethanol-butanol

The utmost important characteristic of such a mixed alcoholwater system is homogeneous and heterogeneous azeotropes formed by water and the highest concentration components of ethanol and butanol,respectively.Therefore,a water (94.38%(mass))-ethanol=3.96 % (mass))-butanol=1.66%(mass)) ternary system is introduced to analyze the separation of aqueous mixed alcohol feed systems as a representative example.

The ternary phase diagram and residue curve maps of the water-ethanol-butanol system at 101.3 kPa are presented in Fig.2(a).The distillation boundary formed by the line connecting the ethanol-water homogeneous azeotropic point and the butanol-water heterogeneous azeotropic point divides the ternary phase diagram into two distillation regions.Since the three components to be separated locate in different distillation regions,the complete separation cannot be achieved with a single column.The feed composition,the sidestream composition,the top and bottom product compositions,and the solubility curves are labeled on the ternary phase diagram.The primary function of the dehydrating column is to obtain high purity water at the bottom of the column,while the purity of the top component is not strictly regulated,so it is reasonable to replace the dehydrating column with a stripping column.Part of the residue curve locates in the two-phase region from the solubility curve,which indicates that the liquid-liquid phase separation phenomenon will occur on the specific column stages in the sidestream distillation column.Therefore,a sidestream distillation column connected to a sidestream decanter can cross the distillation boundary and complete the separation task.In addition,as indicated by the solubility curves at different temperatures,the area of the two-phase region is the largest at 323.15 K.Therefore,there exists an optimal phase separation temperature,which allows butanol to be extracted from the organic phase to the maximum extent.

As shown in Fig.2(b),for the stripper-sidestream decanter process,a mixture(M2)of fresh feedFand decanter aqueous phaseA1enters the stripping column T1,where pure water productB1is extracted from the bottom of the column,and streamD1is obtained from the top of the column.D1is the feed to the sidestream distillation column (T2) and is separated into the butanol productB2and the aqueous ethanolD2.A liquid phaseM1is led into the decanter from the side of the sidestream distillation column T2,where the phase splitting phenomenon occurs,organic phaseO1returns to the next column stage at the extraction location,aqueous phaseA1returns to the stripping column T1.The resultant aqueous ethanol productD1and pure waterB1fall in different distillation regions from butanol productB2,indicating that the addition of the decanter achieves the crossing of distillation boundaries.

2.3.Separation of multicomponent mixed alcohols

The physical properties of aqueous mixed alcohol are similar to the water-ethanol-butanol ternary system owing to the coexistence of homogeneous and heterogeneous azeotropes in both.Hence,the stripper-sidestream decanter process is speculated to be suitable for efficiently separating aqueous mixed alcohols.Herein,only the liquid extraction is considered since the vapor sidestream could not be directly phased and would cause a significant increase in the heat duty of the sidestream distillation column reboiler.The feed enters the stripping column for C1-C8 mixed alcohol dehydration to obtain the C1-C8 alcohol-water mixture at the top and pure water at the bottom.Subsequently,the top product flows into the sidestream distillation column for C1-C8 alcohols cutting,resulting in the C1-C3 alcohol-water mixture as the top product and withdrawing the C4-C8 alcohol-water mixture from the sidestream into the decanter.The organic phase returns to the column,and the dehydrated C4-C8 alcohols are extracted from the bottom of the column.The aqueous phase enters the stripping column for dehydration.

Fig.8.Temperature profiles of (a) stripping column and (b) sidestream distillation column.

Fig.9.Flowsheet of the control strategy.

Fig.10.Controller faceplate of the control strategy.

The temperature dramatically affects the solubility of the C4-C8 alcohols in the aqueous phase.Based on the solubility trend of C4-C8 alcohols in aqueous and organic phases with temperature,330.15 K can be selected as the optimum temperature for the decanter [34].The feed is preheated by pure water at the bottom of the stripping column and then fed into the column.Meanwhile,to reduce the disturbance of the return flow to the stripping column and reduce the energy consumption of the reboiler,the aqueous phase of the decanter is fully heat exchanged with the sidestream of the sidestream distillation column and then returns to the stripping column.

2.4.Optimization of the stripper-sidestream decanter process

In the optimized design of a sidestream distillation column,the location of the extraction is a highly critical design parameter.A single-column process is established in Fig.3(a).The feed is the raw material after preliminary dehydration,and the theoretical stage numbers are initially set to 50 according to the difficulty of separation,and the feed stage is set to 30.The liquid phase composition profiles in the single column are shown in Fig.3(b).The concentration of C1-C3 alcohols shows a monotonic decreasing trend as the column position decreases.After the 34th stage,the concentration of the components is nearly zero.The C4-C8 alcohol concentration maintains an S-shaped increasing trend above the 28th stage but falls precipitously to 31.8%(mass)at the 29th stage and remains stable from the 30th-49th stages.The water content in the column gradually increases as the column stage position decreases,with a jump in water content at the 29th stage and remains at around 64.0% (mass).The purpose of the sidestream extraction is to remove the maximum amount of water from the column and avoid the entrapment of light compositions of C1-C3 alcohols into the stripping column,which causes an increase in energy consumption due to its repeated vaporization.Therefore,it is reasonable to set the sidestream extraction location below the 30th stage(feed stage)in the column.The organic phase return position is the next stage in the sidestream extraction position to maintain the stability of the liquid flow rate in the column.

Sequential iterative search method is widely used to optimize the process [35-39].The optimal values of the decision variable are obtained to achieve the minimum TAC [40].Using TAC minimization as the objective function,the sequential iterative search method is employed for steady-state optimization of the stripper-sidestream decanter process (Fig.4).TheNS,ND1,andNA1are exploited as the inner iteration circle.TheNT1of the stripping column andNT2of the sidestream distillation column are used as the sub-outer iteration and the outer iteration.

Fig.11.Dynamic responses for disturbance in the feed composition in CS1.

Based on the sequential iterative optimization sequence,the optimal sidestream withdrawn location,the theoretical stage number of the stripping column and the sidestream distillation column are determined.With the composition of the feed and C1-C3 alcohol-water azeotropes,the mass flow at the top of the sidestream distillation column is set at 1032.74 kg·h-1.Accordingly,RR2,BR1,andLSare regulated to ensure that the product purity in both columns meets the design specifications.Fig.5(a)shows the effect ofNSon the reboiler duty of the sidestream distillation column.The reboiler duty reaches a minimum at the 32nd stage.Therefore,the optimal sidestream position is the 32nd stage.Fig.5(b) shows the effect ofNT1andNT2on TAC,and there exists an optimalNT1making the lowest total TAC whenNT2is constant.The TAC is lowest whenNT1=47 andNT2=45.Fig.6 displays the as-proposed optimal stripper-sidestream decanter process.

2.5.Conventional three-column sequence

Fig.7 shows a conventional three-column process which is established and optimized with the same feed.The extraction streams data are consistent with those of the stripper-sidestream decanter process [34].

2.6.Comparison

The comparative data of the two processes are listed in Table 1.In the fraction cutting process,the overall process achieves precise separation of the alcohol-water mixture.However,the overall process of the traditional process is complex—the multiple distillations of C1-C8 lead to high process energy consumption.Moreover,the results show that the stripper-sidestream decanter sequence can reduce TAC by 17.00% and save energy by 21.27%compared with the traditional three-column process.When replacing the dehydrating column with a stripping column,the absence of the rectifying section brings a weaker column separation capacity.More water enters the sidestream distillation column (alcohol separation column),making the energy consumption of the reboiler of the stripping column increase by 0.82 %compared with that of the dehydrating column.However,the vapor material at the top of the T1column enters the next column without condensing,which not only reduces the use of cooling water by eliminating a condenser but also avoids the re-vaporization of the liquid phase material and saves significantly on the reboiler energy consumption of the sidestream distillation column.By analyzing the liquid flow components of the whole column,a liquid flow is led into the decanter at the column stage where the phase split occurs,the organic phase is returned to the column,and the aqueous phase re-enters the stripper.The separation of water and slightly soluble C4-C8 alcohols in the column through the decanter reduces the difficulty of separation in the sidestream distillation column and avoids using heavy alcohol distillation columns ofmixed alcohols,and reduces energy consumption and equipment costs.

Table 1 Comparison of the stripper-sidestream decanter sequence and three-column sequence.

Table 2 Controller tuning parameters in CS1.

Table 3 Controller tuning parameters in CS2.

Fig.12.Dynamic responses for disturbance in the feed flow rate in CS1.

3.Dynamic Simulation and Basic control Structure

The control scheme is established for the stripper-sidestream decanter process.Feed composition disturbances and flow rate disturbances are introduced to test the performance.The temperature profiles of the stripping column and sidestream distillation column are shown in Fig.8.The sensitive tray is selected according to the slope criterion proposed by Luyben [41].

The first stage of the stripping column and the 17th and 35th stages of the sidestream distillation column are finally determined to be the sensitive tray.Dynamic performance is tested by introducing a ±20% feed composition and flow rate disturbances at 0.2 h.Ethanol,the highest concentration of the alcohol fraction,is selected as the component of the composition disturbances.

To ensure that the materials can be properly transported throughout the piping system,the dynamic simulation process requires the addition of necessary components such as pumps,compressors,and control valves to make the process a pressuredriven one.The addition of the compressor creates a slight variation between some of the dynamic control parameters and the steady-state process parameters.

3.1.Control scheme 1 (CS1)

The control loop is listed below:

(1) The flow rate of the fresh feed is controlled by the valve opening of the flow controller.

(3) The liquid levels of the sump of the sidestream distillation and stripping columns are controlled by the bottoms rate of the columns.

(4)The aqueous phase and organic phase levels of the decanter are controlled by their corresponding aqueous phase rate and organic phase rate,respectively.

(5) The temperature of the sidestream entering the decanter is maintained stable by the heat exchanger.

(6)The pressure at the top of the sidestream distillation column is controlled by the heat removal of the condenser.

(7)The pressure at the top of the stripping column is controlled by the compressor power,which is in cascade control with the FDC(the distillate rate controller) because its set value signal comes from the output of the top pressure controller.

(8)The reboiler duty(QR1)of the stripping column is controlled proportionally to its feed rate setting.QR1/F1is in cascade control because the set value signal of controllerQR1/Fcomes from the output of temperature controller TC1 which control the temperature of the 1st stage of the stripping column (T1).

(9)The temperature of the 17th stage of the sidestream distillation column (T17) is controlled by the reflux ratio of the column.

(10)The temperature of the 35th stage of the sidestream distillation column (T35) is controlled by the reboiler duty of the column.

15. Lose them again: That the scenario81 is repeated could be read as the child s need to master not only the internal world, but also the external-as much scarier place sometimes, with circumstances and events beyond one s control. Return to place in story.

Fig.13.Dynamic responses for disturbance in the feed composition in CS2.

(11)The sidestream extraction flow rate is controlled by a lowselector with control signals from Sum31-32,a summation controller,andQR2/SIDE,a proportional controller.The output signal of Sum31-32 is the sum of the liquid phase flow rates of the 31st and 32nd stages of the column.QR2/SIDE is the ratio of the reboiler duty to the sidestream extraction.

The control strategy of the whole process is shown in Fig.9.The low-selector scheme is used for simultaneous and timely control of the sidestream extraction flow rate to avoid the impact of unreasonable set values of the sidestream extraction flow rate on the column.Due to the presence of measurement and actuator lags,a 1 min deadtime is inserted in the temperature control loops.Relay-feedback tests are performed on the three temperature control loops to obtain gain constants and integration time constants using the Tyreus-Luyben tuning method,and the tuning results are shown in Table 2.Considering that the sidestream extraction cooler can respond quickly to changes in the temperature of the flow,it can be assumed that its temperature control loop has no dead time,and the TC-DEC temperature controller does not need to be tuned.The control faceplate of the control strategy is shown in Fig.10.

The dynamic responses for feed composition disturbance are given in Fig.11.T1,T17,and T35 approached the set values after 5 h.However,in+20%composition disturbance,there is an intensive small periodic oscillation around the set value after 5 h,and for -20% composition disturbance,there is an intermittent step after 5 h.The mass fraction of butanol (ωD2C4) in the C1-C3 alcohol-water mixture and the mass fraction of 1-propanol (ωB2C3)in the C4-C8 alcohol-water mixture and water (ωB2H2O) in the C4-C8 alcohol-water mixture is stabilized around the set values after 5 h.The mass fraction of water in the pure water product(ωB1H2O) is basically non-volatile in the face of +20% composition disturbance.As for -20% composition disturbance,the system reaches a new steady-state after fluctuation,with some deviation from the original set point.

Fig.12 gives the dynamic responses for feed flow rate disturbance.The results are almost identical to those of feed composition disturbance.The main difference is that T1,T17,and T35 do not produce intermittent steps after reaching the steady-state after 5 h at -20% feed flow rate disturbance.Meanwhile,ωB1H2Ois non-volatile.At+20 % feed flow rate disturbance,ωB2H2Oand ωB2C3return to the original set point after a large transient deviation.

The above result indicates that CS1 fails to control the+20%composition and flow rate disturbance effectively.The temperature-sensitive stage position controlled by the TC35 controller is close to the sidestream extraction location,and its temperature is easily disturbed by fluctuations in the extraction flow rate.Meanwhile,the control deficiency of CS1 is that the PI controller cannot regulateQRin time to respond to temperature changes due to its hysteresis characteristics.A more robust control strategy will be investigated in the next section to solve this problem.

3.2.Control scheme 2 (CS2)

Fig.14.Dynamic responses for disturbance in the feed flow rate in CS2.

Differential control responds to the input quantity variation and can sensitively sense fluctuations,allowing the controller to respond as early as possible,thus improving the response speed and dynamic performance of the system.This section builds the control structure CS2 based on the CS1 control structure.Differential control is introduced in the CS2 control structure to suppress parameter perturbations in advance.The PID tuning of the TC35 controller is tuned using the Ziegler-Nichols method,and TC1 and TC17 are retuned using the Tyreus-Luyben method.The specific tuning parameters of the three controllers are shown in Table 3.It can be seen that after the introduction of differential control,the proportional gain of TC35 increases and the integration time decreases notably,which significantly improves the control efficiency of the temperature controller;after returning,the proportional gain of TC1 increases,the integration time of TC17 decreases slightly,and the overall process control structure is more responsive and efficient.

Fig.13 displays the dynamic response results of the feed composition disturbance.Facing the +20% feed composition disturbance,T1,T17 and T35 will no longer oscillate periodically and can play a good control effect.Under -20% feed composition disturbance,T1,T17,T35,ωB2H2O,and sidestream distillation column reboiler dutyQR2exhibits step changes at 9 and 19.5 h due to lateral stream extraction close to 32nd stage liquid phase flow rates.However,the differential controller can quickly return to the steady-state point because of its highly sensitive nature,resulting in a satisfying overall dynamic control.

Fig.14 shows the dynamic response results with the disturbance of a 20%feed flow rate.For the+20%feed flow disturbance,T17 and T35 returned to the original set point after 5 h,eliminating the cyclic oscillation.The transient deviation of 1-propanol and water concentrations in the C4-C8 alcohol-water mixture became significantly smaller and stabilized near the setpoint after 5 h.It is shown that the introduction of the differential controller makes it possible to resist flow rate disturbances of +20%.

In summary,the CS2 has excellent dynamic controllability for both composition and flow rate disturbances.The controlled temperature point can quickly return to the set point,and the product concentration can be stabilized near the original set point to achieve excellent separation of the C1-C3 alcohol-water mixture,C4-C8 alcohols and water.Therefore,although the process of the stripper-sidestream decanter is slightly complicated,the proposed control structure shows robust control,which also lays the foundation for further applications of the process.

4.Conclusions

In this study,a sripper-sidestream decanter sequence is established and optimized to achieve the energy-efficient separation of mixed alcohols from Fischer-Tropsch aqueous by-product and to accurately obtain C1-C3 alcohol-water mixtures,dehydrated C4-C8 alcohols and pure water.In this sequence,the strippingcolumn reduces the redistillation of C1-C8 alcohol-water mixture.Besides,the sidestream distillation column saves energy by using the decanter to take on some of the dehydration tasks.The stripper-sidestream decanter sequence is economically optimized with the TAC minimization.The results show that,compared with the conventional three-column separation process,The TAC is reduced by 17.00%,and the energy consumption is reduced by 21.27%.Meanwhile,the proposed process control structure is investigated.The control scheme CS1 and CS2 of process are established.The results show that CS1 is resistant to -20% feed composition and flow rate disturbances and not resistant to +20%disturbance.The introduction of the differential controller in CS2 made it possible to resist the disturbance of the sidestream extraction to the sidestream distillation column.The overall process exhibited excellent dynamic controllability.

Nomenclature

BR boilup ratio

FCI fixed capital investment

irfixed capital recovery rate

KCgain

LSwithdrawal flow rate of sidestream

NA1feed stage of the aqueous phase entering the stripping column

ND1feed stage of the sidestream distillation column

NSwithdrawal location of sidestream

NTtheoretical stage number

OC operation cost

PI proportion integration

PID proportion integration differentiation

QRreboiler duty

RR reflux ratio

TAC total annual cost

τDderivative time

τIintegral time

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

Dedicated to well-beloved Professor Guocong Yu (K.T.Yu) and acknowledged the encouragement from him four years ago.The respected Professor Guocong Yu has always been serving as a role model.As a postgraduate in Chemical Engineering at Tianjin University,I will certainly try my best to look forward to the future and realize my dream to meet his expectations.