Coupled fluidized bed reactor for pyridine synthesis - Industrial

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Kinetics, Catalysis, and Reaction Engineering

Coupled fluidized bed reactor for pyridine synthesis Shuaishuai Zhou, Shanhe Liu, Yongfei Wei, Xihong Li, Chunxi Lu, mengxi liu, and Xiaotao Bi Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b01510 • Publication Date (Web): 15 May 2018 Downloaded from http://pubs.acs.org on May 18, 2018

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Industrial & Engineering Chemistry Research

Coupled fluidized bed reactor for pyridine synthesis Shuaishuai Zhou1, Shanhe Liu2, Yongfei Wei2, Xihong Li2, Chunxi Lu1*, Mengxi Liu1*, Xiaotao Bi3 1. State Key Laboratory of Heavy Oil Processing, China University of Petroleum, Beijing 102249, PR China. 2. Costar Biochemical, Anhui 243100, PR China. 3. Department of Chemical and Biological Engineering, University of British Columbia, Vancouver BCV6T1Z3, Canada *Corresponding author. Tel./fax: +86 10 89733803. E-mail address: [email protected], [email protected]

Abstract：To eliminate the feeding device coking deposit in commercial pyridine synthesis reactor, a coupled fluidized bed reactor is proposed. The coupled reactor is composed of a feeding zone (FZ), a riser reaction zone (RRZ) and a fluidized bed reaction zone (FRZ). During 15 days of continuous operation, there is no coke deposit in the feeding device. Experimental results show that the product yield reaches as high as 75 %. The selectivity is around 2.5 and is higher than that in commercial reactors, which is around 2.2. Core-annulus model and Dispersion model are proposed to model different zones of this coupled reactor, and the predicted average deviation from experimental data is 12 %. The prediction results show that the main reactions take place at FZ and FRZ. RRZ contributes the least conversion, because of the limit by the mass transfer between the core and annulus. Key words: Coupled, Fluidized bed reactor, Pyridine synthesis, Experiment, Modeling

1 INTRODUCTION Pyridine and 3-picoline have been widely used for synthesis of pharmaceuticals and agrochemicals, benefiting from high chemical reactivity as well as biological activity 1. Traditionally, pyridine was derived from coal tar and suffered from high sulfur content. In 1920s, high market demand of pyridine encouraged the development of new technologies for pyridine synthesis. Nowadays, most pyridine is produced from the Chichibabin condensation process 1. Research on pyridine synthesis has been rare and mainly focused on: catalysts, reaction mechanisms, reaction kinetics and reactors. The development of catalyst can be divided into two stages 1, which are silica-alumina catalysts and zeolite catalysts. Until now the most widely used zeolite catalysts are HZSM-5 catalysts, benefiting from the design of an appropriate shape selective system. Because of multiple side reactions occurring in pyridine synthesis, the main reaction mechanism of the process has not been well understood. It has only been confirmed that during the reaction process, imide intermediates is first produced, which is the key component to form pyridine kinetics of pyridine synthesis reaction.

Reddy 7

2-6.

There have been few reports on the reaction

carried out experiments in a 20 mm ID Pyrex reactor using

HZSM-5 catalysts, and their results suggested that the reaction rate is only determined by reactor operating conditions and acetaldehydes concentration. Fluidized bed reactor is favored for pyridine synthesis. Compared with fixed bed reactor, fluidized bed reactor has advantages of high heat and mass transfer rates, good gas and catalyst contact and capable of continuous catalysts regeneration. Besides, previous studies reported higher product yield of fluidized bed reactors than fixed bed reactors for pyridine synthesis

8-11.

Feitler et al.

11

produced pyridine in both fixed

bed and fluidized bed reactors. Over a large range of reaction temperatures from 723 K to 773 K, the

*Corresponding author. Tel./fax: +861089733803. E-mail address: [email protected]. 1 Environment ACS Paragon Plus

Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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fluidized bed reactor offered higher product yield than the fixed bed reactor. Due to those advantages of fluidized bed reactors, some commercial scale pyridine synthesis units have been incorporated with fluidized catalyst beds. The gas-solid flow patterns in fluidized bed reactors ranged from bubbling fluidization to fast fluidization. Usually, the flow pattern in commercial reactors for pyridine synthesis is bubbling fluidization. A problem that results in frequent shut down of commercial reactors is the coking deposit inside the feeding device. Coke is formed due to the high temperature around the feeding device reactor. To solve the problem, Zhou

12

12,13

, which is installed inside the

proposed to use a fast fluidized bed reactor to replace the bubbling

fluidized bed reactor for pyridine synthesis. By changing the flow pattern from bubbling to fast fluidization, the diameter of the reactor decreases, and the feeding device could be installed outside the reactor to avoid high temperature environment. Zhou

12

synthesized pyridine in a fast fluidized bed reactor with a 14 mm

ID and 3500 mm tall reactor, and the highest product yield is 65 %. Compared to the product yield of commercial reactors, which is 73 %, the product yield decrease probably results from the short gas-solid contact time or slow reaction rate. To improve the performance of fast fluidized bed reactor, a coupled reactor is proposed in this paper. The coupled reactor is composed of a feeding zone (FZ) with novel feeding scheme, a riser reaction zone (RRZ) to increase the selectivity of pyridine over 3-picoline and a fluidized bed reaction zone (FRZ) for enough gas-solid contact time. The details of this coupled reactor can be found in the experimental apparatus section. For coupled reactor, previous studies show that the flow patterns range from bubbling to fast fluidization in different sections along the bed height

14,15.

For feeding zone (FZ), the gas velocity is lower than that in

riser reaction zone (RRZ) above it, and the solid concentration is higher than that in RRZ. Previous studies showed that non-uniform flow distribution in the FZ, which is caused by unsymmetrical catalyst feed, can be eased by setting an enlarged section at the bottom of the riser. Liu

16

investigated the FZ of a riser by

experiment, and they found that the non-uniform flow problem is reduced and gas-solid contact is enhanced. The riser reactor with FZ gives better reaction performance has been demonstrated in Fluid Catalytic Cracking (FCC) process 17-20. In this paper, the arrangement of FZ mainly based on the following considerations, which are to reduce the non-uniform flow phenomenon in the feed section and to eliminate the coking problem in the feeding device. In the RRZ, gas back mixing decreases. Besides, the particle residence time also decreases, which brings high gas-catalysts contact efficiency

21-23.

The flow pattern in RRZ is fast fluidization where the solid

concentration axial distribution exhibits from exponential shape to C shape depending on the reactor exit configuration 24-26. Usually, a dense bed is formed at the bottom, with dense bed height determined by the properties of the gas-solid system and operating conditions

27-31.

Based on experiment, Li

32

proposed a

semi-empirical model to predict the solid concentration axial distribution in fast fluidized bed reactors 33. Along the radial direction in the fast fluidized bed reactor, there exists a core-annulus structure 34,35. In the core region, the solid concentration is low, and the gas concentration is high. While, in the annulus region, the structure is opposite. Usually, the mass transfer between the core and annulus, which is similar as the

2

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mass transfer between dense phase and bubble phase in bubbling beds, significantly affects the reactor performance. To quantify the mass transfer between the core and annulus, some useful correlations

36-38

have been developed based on experimental results. In the upper FRZ, the gas velocity decreases to around 0.5 m/s. For the gas-solid system investigated in this paper, the flow pattern is turbulent fluidization 39. By adjusting the height of the dense bed in FRZ, this section serves either as a disengagement zone or as a reactor. Details on this zone will be presented in the experiment apparatus section. Numerous works have been reported on the turbulent fluidization and fast fluidization individually. Few reports are found on the coupling of these flow patterns in one reactor. Wang 15 investigated the solid phase flow hydrodynamics in a coupled fluidized bed reactor. The experiment apparatus employed in Wang’s work is made of plexiglass, and the experimental results are obtained under cold operating conditions. They found that the axial solid concentration distribution in the middle RRZ section is typically of C shape, and the influence of the dense bed in FRZ on RRZ is mainly reflected by the solid concentration in RRZ. With the increase of the dense bed height in FRZ, the solid concentration in RRZ is increased, especially in the region near the riser exit. Based on experimental results, Wang proposed some empirical correlations to predict the solid concentration axial distribution in the RRZ. This coupled reactor has been successfully applied in the FCC gasoline olefins reduction process. While, no report is found on the application of the coupled reactor in the pyridine synthesis process. This paper aims at investigating the performance of coupled reactors for pyridine synthesis.

2 EXPERIMENT Figure 1 shows the schematic diagram of the experimental apparatus (Costar Biochemical, Ltd). As illustrated in Figure 1, the coupled reactor composes of FZ (feeding zone), RRZ (riser reaction zone) and FRZ (fluidized bed reaction zone). The height of the FZ is 1.0 m, and the ID is 140 mm. For RRZ, the total height is 13.5 m, and the ID is 84 mm. The FRZ is 5.74 m in height and has an ID of 311 mm. This coupled reactor is made of stainless steel. Formaldehydes and acetaldehydes feedstock flow were preheated to 373 K, with the flow rate measured and adjusted by a regulating system, composed of flow meter, a valve and a control device (Costar Biochemical, Ltd). They were mixed with ammonia, which was regulated by another regulating system. The mixture of feedstock was introduced into the FZ by a nozzle, mixing with hot regenerated catalysts from an adjacent commercial regenerator and vaporizing. The vapor and catalysts flowed upwards in the coupled reactor and were separated by a cyclone at the top. The gas product was introduced to the commercial reactor by products pipeline. To analyze the products composition, a sampling pipeline is installed on the products pipeline, with a side stream of the products being withdrawn. The withdrawn products were condensed and then collected into a glass bottle, with the liquid products composition being determined by a gas phase chromatograph (GC7890II, Shanghai Tianmei Keji Instrument Co., Ltd., China). The spent catalysts were stripped by water steam and then returned to the industrial reactor. After regeneration, the spent catalyst was conveyed to coupled fluidized bed reactor through catalysts circulation 3



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pipe. The initial acetaldehydes concentration cA0 was calculated by eq (1) to eq (4).

cA0 =

nA0 V

(1)

where nA0 is the mole flow rate of acetaldehydes, which was obtained by multiplying the mass flow rate of aldehydes flow and its corresponding composition. V is the volumetric flow rate of both aldehydes flow and ammonia at reaction conditions, which was derived based on ideal gas law.

PV = nRT

(2)

where n is the mole flow rates of both aldehydes and ammonia and can be calculated by the following equation.

n = nA0 / xA0 + nm

(3)

where xA0 is the mole fraction of acetaldehydes in aldehydes flow, nm is the mole flow rate of ammonia and was calculated by

nm = (1/ af + 1)nA0 / am

(4)

where ηaf is acetaldehydes to formaldehydes ratio by mole (ATFR), ηam is ammonia to aldehydes molar ratio (ATAR), which were obtained by the composition of both aldehydes and ammonia before experiment. Reddy 7 suggested that the volume change in a small reactor can be neglected, therefore the gas velocity was assumed to be constant in the coupled fluidized bed reactor. The superficial gas velocity in the entrance of the coupled fluidized bed reactor is calculated by eq (5).

ug =

4V  D2

(5)

where D is the internal diameter of the coupled fluidized bed reactor. The reaction temperature and pressure were measured by thermocouples and pressure gauges, all thermocouples and pressure gauges were linked to a computer via a data acquisition system. The positions of the thermocouples and pressure gauges are depicted in Figure 2. As shown in Figure 2, there were two temperature taps, with one located in the FZ (2-1, T1) and the other one in the FRZ (2-2, T2). By neglecting the wall friction and acceleration/deceleration, the average solid concentration in RRZ is calculated from measured pressure drop by eq (6).

 fs =

P1−1 − P1−2 ( z1−2 − z1−1 ) g

(6)

The solid concentration of the dense bed in FRZ is calculated by eq (7).

bubs =

P1−3 − P1−4 ( z1−3 − z1−2 ) g

The height of the dense bed in FRZ is calculated by eq (8).

4


(7)

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hd =

P1−3 − P1−5 bubs g

(8)

It was difficult to measure the catalysts circulation rates directly. Alternatively, the heat balance method was used to calculate the catalysts circulation rates. Figure 3 shows the energy balance volume used for catalysts circulation rate calculation. Based on heat balance of the control volume under steady state, eq (9) is obtained. By solving eq (9), the catalysts circulation rate could be obtained.

 mi H i +

 Dtur 2Gs 4

L(Tca − To ) =  mo H o

(9)

Since only a small fraction of products was collected, it is impossible to calculate the product yield by dividing the mass of pyridine in product with the mass of aldehydes in the feedstock. In this paper, the product yield is calculated based on unit feedstock, with the details given as following. Firstly, the aldehydes weight in unit feedstock is calculated by eq (10). Then the pyridine weight in unit product is calculated by eq (11). At last, the product yield of pyridine is calculated by eq (12).

mAfe = M A / M fe mApr = (

cDpr MD

 2M A +

cEpr ME

(10)

 2M A ) / (1 − cam )

(11)

y = mAfe / mApr

(12)

where mAfe is the mass of acetaldehyde in unit feedstocks, Mfe is the mass of feedstock, mApr is

the mass of acetaldehyde in unit products, cDpr is the concentration of pyridine in product, cEpr is the 3-picoline concentration in products, cam is the ammonia concentration in product, y is the product yield of pyridine and 3-picolines.

3. MODELING In this paper, a coupled reactor model is proposed to investigate performance of this coupled reactor. The details of the model are as following. 3.1 Reaction Kinetics Up to now, the research on reaction kinetics of pyridine synthesis has been rare. Reddy

7

proposed a

reaction mechanism based on parallel series reactions. As shown in Figure 4, the feedstock (mixture of acetaldehyde, formaldehyde and ammonia) first produce pyridine and 3-picoline, then a part of pyridine and 3-picoline form coke at different rate. The main reactions are given by eq (13) and eq (14). The byproducts in pyridine synthesis process are 2-picoline, 4-picoline, methylamine, dimethylamine and trimethylamine, which can be minimized by varying the composition of the feed-stock and the ratios of

5



Page 6 of 24

mixture.

2CH3CHO+HCHO+NH3 ⎯⎯ → C5H5 N+3H2O+H2

(13)

2CH3CHO+2HCHO+NH3 ⎯⎯ → C6H7 N+4H2O

(14)

Based on the above mechanism and experimental results, Reddy

7

proposed the following reaction rate

expressions, with the fitted reaction constants and activation energy listed in Table S1.

−rA = (k1 + k2 )CA 2a

(15)

rD = (k1CA 2 − k3CD )a

(16)

rE = (k2CA 2 − k4CE )a

(17)

where a is the reactivity of catalysts, and the value is unity for fresh catalysts. In this paper, the reactions and the catalysts are the same as used in Reddy’s experiments, and the above reaction kinetics is adopted. 3.2 Reactor Model For this coupled reactor, the gas-solid flow pattern varies from turbulent fluidization to fast fluidized bed 39 along the reactor axial direction. As shows in Table 1, for FZ and RRZ, the flow pattern is fast fluidization. In FRZ, the flow pattern is turbulent fluidization. Table 1 The flow patterns in different sections of the coupled reactor ug (m/s)

flow pattern

FZ

1.7 ~ 2.6

fast fluidization

RRZ

4.8 ~ 7.1

fast fluidization

FRZ

0.52 ~ 0.72

turbulent fluidization

For different flow patterns, various reactor models are proposed. For fast fluidization, the reactor could be modeled by the core-annulus model 34-35. For turbulent fluidization, various models are proposed 40-41, and these models could be divided into two main categories, which are CSTR in series model and dispersion model. In this paper, different reactor models are chosen for different sections of the coupled reactor, with the details given as following. 3.2.1 Reactor Model for FRZ The flow pattern in the FRZ is turbulent fluidization. For turbulent fluidized bed reactors, the widely used models are CSTR in series model and dispersion model. As shows in Figure 5, CSTR in series model uses ideal mixed reactors to simulate the gas residence time distribution (RTD)

42.

While for the dispersion

model, gas back mixing is represented by an axial dispersion diffusion coefficient, which can be estimated by the correlation proposed by Foka 43.

Pe = 0.071Ar 0.32 (d p / Dtur ) −0.4

(18)

Levenspiel 44 pointed out that for flow patterns not far from plug flow, the dispersion model should be used. The Pe number was found to be around 3 under our experimental operating conditions, so the dispersion model is adopted in this study. Based on gas phase mass balance, eq (19) can be derived for gas phase axial dispersion flow 45.

6


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ug

dc d 2c − Dzg 2 + r = 0 dz dz

(19)

By combing eq (19) with reaction kinetics, the following equations can be obtained for species A, D, E and F, which form the model for FRZ.

dcA d 2cA − Dzg − (k1 + k2 )c A 2  s = 0 2 dz dz dc d 2 cD u gtur D − Dzg + (k1c A 2 − k3cD )  s = 0 dz dz 2 dc d 2 cE u gtur E − Dzg + ( k 2 c A 2 − k 4 cE )  c = 0 2 dz dz dc d 2 cF u gtur F − Dzg + ( k3 c D + k 4 c E )  s = 0 dz dz 2 ug

(20)

To solve the above model equations, the catalyst concentration distribution along the axial direction is required. In this paper, the correlation proposed by Lu

46

is modified and adopted. Lu used three different

gas-solid systems to investigate the catalyst concentration axial distribution for turbulent fluidized bed reactor by experiment. Based on experimental results, Lu proposed three different correlations to calculate the catalysts concentration, which are given in Table 2. Table 2 Correlations for the particle concentration calculation

ug (m/s)

Particle Properties

Correlation

0.535~1.328

ρpb=612 kg/m3，dp=63.3 μm

ρs=-77.62z-141.56ugtur+503.39

0.536~1.315


ρs=-89.95z-126.49ugtur+502.64

0.826~1.519


ρs=101.62ugturz-387.15-252.94ugtur+925.3

Table 2 shows that the correlations proposed by Lu did not take the influence of gas-solid properties into consideration, which in turn generated three correlations for three different gas-solid systems. For simplification, in this study, the following correlation is proposed.

 s = Ar k (k2ugtur + k3 z + k4 ) 1

(21)

By fitting Lu’s experiment data, the values of the parameters k1, k2, k3 and k4 are obtained, which are k1=1.013, k2=-12.626, k3=-10.015, k4=48.936. Figure 6 is the comparison between the prediction result by modified correlation and experimental data, with an average prediction deviation of 7 %. 3.2.2 Reactor Model for RRZ The flow pattern in FZ and RRZ is fast fluidization. Core-annulus model is chosen to model this section. Figure 7 shows the basic concept of core-annulus model, which divides the cross-sectional area into a core and an annulus region. In the core, the solid concentration is low, and the gas phase could be treated as plug flow. While in the annulus, the solid flow downward and the back-mixing is severe, and the gas phase is treated as mixed flow. Based on mass balance in the core region and annulus region, the following equations are obtained. Core region:

7



dcc 4kac + (cc − ca ) + r = 0 dz Dr  g1/2

u gc

Page 8 of 24

(22)

Annulus region:

dca 4ka  c1/2 (ca − cc ) + r = 0 u ga + dz Dr (1 −  c ) Marmo

47

and Kagawa

48

(23)

assumed the gas velocity in the annulus as zero. In this paper, the same

assumption is used. By combing equation (22) and equation (23) with the reaction kinetics, the following equations could be obtained.

caA =

−m + m 2 + 4mnccA 2n

(m =

4kac  c1/2 , n = (k1 + k2 )  sa ) D f (1 −  c )

dccA 4kac = (− (ccA − caA ) − (k1 + k2 )ccA 2  sc ) / u gc 1/2 dz D f c caD =

mccD − k1caA 2  sa m − k3  sa

dccD 4kac (ccD − caD ) + (k1ccA 2 − k3ccD )  sc ) / u gc = (− 1/2 dz D f c caE =

(24)

mccE − k2 caA 2  sa m − k4  sa

dccE 4kac = (− (ccE − caE ) + (k2 ccA 2 − k4 ccE )  sc ) / u gc dz D f  c1/2 dccF = (k3ccD + k4 ccE )  sc / u gc dz where Фc is ratio of the core area to cross sectional area and is calculated by the correlation proposed by Patience 49.

c =

1 1 + 1.1Fr (u ps / u gf ) 0.083 Fr

(25)

The solid concentration radial distribution could be treated as follows. The solid fraction in the core region is assumed as a constant, which is equal to 0.005

50-52.

The solid concentration in the annulus could be

calculated by mass balance in eq (26).

 c c + (1 −  c ) an =  av

(26)

where the axial solid concentration εav is calculated by Li-Kwauk model 53,54, which is shows as following.

ln(

 av −  bt 1 ) = − ( z − zi )  '−  av Z0

The mass transfer coefficient is calculated by the correlation proposed by Patience 49,

8


(27)

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kac D f Dg  g

 g1/2 = 0.25(

 g 0.5  g ugf D f 0.75 ) ( ) [Gs / (  pug )]0.25 1/2 Dg  g  g (c )

(28)

where Dg is the gas diffusion coefficient and is calculated by the correlation proposed by Fuller 55.

1 1 + MA MB DAB = P[( vA )1/3 + ( vB )1/3 ]2 0.001T 1.75

(29)

where DAB is the binary diffusion coefficient of gases A and B, ∑vA and ∑vB are the molecular diffusion volume of A and B respectively. Equations (24) to (29) form the reactor model for the RRZ.

9



Page 10 of 24

4 EXPERIMENT RESULTS 4.1 Results of this pilot scale reactor The experimental operation conditions are listed in Table 3. The physical properties of gases are obtained from ASPEN Plus. The experimental apparatus successfully run for 15 days and then stopped for safety check. There was no coke formation on the feed nozzle, which indicates the coking problem in the feed introduction device has been partly eliminated. Table 3. Variation range of operating parameters Parameters

Variation range

cA0 (mol/L) T1-1 (K) T1-2 (K) P2-1 (kPa) Gs (kg/m2/s) hd (m) dp (mm) ρp (kg/m3) ρg (kg/m3) μg (Pa.s)

0.003~0.009 711~757 700~734 180~190 10~80 0~2.2 70 2400 0.71 2.56x10-5

Gs: catalysts circulation rate based on the RRZ cross sectional area

During the experiments, the values of ATFR and ATAR were in unity, ratios widely used in pyridine bases industries. Table 4 and Table S2 showed the experimental production yield. Table 4. Selectivity and Product yield under different operating conditions N

cA0

T1-1

T1-2

Gs

hd (m)

Selectivity

yield (%)

(mol/L)

(K)

(K)

(kg/m2/s)

1

0.003

729

708

25

1.3

10.2

25

5

0.007

733

701

47

1.4

4.2

52

7

0.008

741

729

56

1.2

3.0

54

10

0.008

742

708

57

0

2.6

41

14

0.007

725

714

36

1.4

2.9

59

15

0.007

730

714

36

2.2

3.3

58

18

0.007

738

727

72

1.1

2.2

72

20

0.007

711

701

58

2.0

3.3

58

From Table 4 and Table S2, it is seen that the selectivity of pyridine over 3-picoline varies with production yield, and the selectivity value is greater than 2.2. The selectivity under the highest production yield operating conditions is 2.6, and is greater than that from commercial reactors, which is around 2. The increase in selectivity mainly results from the increased gas velocities in this coupled reactor, which shortens the residence time of the catalyst and increases its reactivity 12. The highest product yield is 75 %, which is similar as that of commercial reactor, which is 73 %. This indicated that this coupled reactor has a better performance than single fast fluidized bed reactor in terms of production yield. 10


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From experimental data 1 and 5, it is seen that the increase in initial acetaldehyde concentration raises the product yield, which agrees with the model prediction result of Zhou 12. However, for commercial reactors, the feedstocks are fixed, it is impossible to change the initial acetaldehyde concentration. The influence of initial aldehyde concentration on product yield is thus only of interest for lab scale experiment. Compare experimental data 7 and 10, the product yield is seen to vary from 54 % to 41 % when the reaction condition changes from Hb=1.2 to Hb=0. For Hb=0, the dense bed in FRZ disappears, the reactions take place mostly in FZ and RRZ. For Hb=1.2, an additional reaction region-bubbling fluidization section forms, and this region contributes 13 percent conversion. This result indicates that the product yield decrease observed by Zhou may result from the limitation of gas-solids contact time under certain reaction conditions. From experimental data 14 to 15, the dense bed height in the FRZ increases by 0.8 m, while the production yield keeps almost the same. This indicates that the increase of residence time does not always increases the product yield. The reason is that, with the increase of residence time, the product yield first increases and then decreases due to over-reaction 12. Therefore, there must exist a suitable residence time, which could be determined from experiment. Experimental data 18 and 20 show the influence of reaction temperature on the product yield. It shows that increase of reaction temperature will increase the product yield. 4.2 Comparison between the pilot scale reactor and commercial reactor Table 5 shows the comparison between the pilot scale reactor and commercial reactor. As shows in Table 5, compared with the commercial reactor, this pilot scale reactor has higher product selectivity and lower spent catalysts coke contention with similar product yield. Table 5. Comparison between the pilot scale reactor and commercial reactor performance Product yield (%) Product selectivity Coke content in spent catalysts (w%)

Pilot Scale Reactor

Commercial Reactor

75 Above 2.5 0.92

73 2.2 1.46

The selectivity is determined by reaction kinetics and can be modified through adjusting the flow hydrodynamics. Figure 8 shows the selectivity variation with reactants residence time under typical commercial reaction condition. Figure 8 indicates that the selectivity increases with the increase of reactants residence time. However, to optimize the product yield, the reaction should be terminated before reaching the over-reaction zone 1. Under this reaction condition, the best reactants residence time is 3.1 s. In that case, the reactants residence time should be kept at this value. However, due to the back-mixing, the reactants residence time of real reactors is a distribution rather than a unique value. To narrow down the reactants residence time distribution, the back-mixing should be minimized. For this coupled reactor, the back-mixing in different reaction zones are varying owing to different gas-solid flow patterns. Since the back-mixing of turbulent fluidization and fast fluidization is less than that in bubbling fluidization, the coupled reactor has less back-mixing than the commercial reactor, which results in a higher selectivity of the products. Table 5 shows that the spent catalysts coke contention is higher in the commercial reactor. The spent catalysts coke contention is closely related to the catalysts residence time 12, with the same reaction conditions, the spent catalysts coke contention increases with the catalysts residence time. For the fast fluidization, the catalysts residence time is around 10 s 12. While, for the bubbling fluidization, the catalysts residence time is around 1 h. For this pilot scale reactor, the flow patterns are fast fluidization and turbulent 11



fluidization, and the catalysts residence time is less than that of the bubbling fluidization commercial reactors, which decreases the spent catalysts coke contention.

5 MODELING RESULTS Eqs. (18) to (29) formed the reactor model for this coupled reactor. Comparison between the model prediction and experimental data of this coupled reactor is illustrated in Figure 9. The average prediction deviation is 12 %. Figure 10 shows the gas concentration variation along the axial direction of this coupled reactor, which indicates that the main conversion takes place in FZ and FRZ. The conversion in RRZ only accounts for 10 percent of the total product yield. The superficial gas velocity in RRZ is 6.1 m/s, and the length of the reaction zone is 13.5 m, then the residence time is 2.2 s. For FZ, the superficial gas velocity is 1.5 m/s, and the length of the reaction zone is 1 m. In that case, the gas residence time in the FZ is 0.67 s. From the residence time point of view, the conversion in the RRZ should be higher than that in the FZ. However, the model prediction result gives the opposite conclusion. This indicates the superficial reaction rate in RRZ is low. Damokler (Da) number is used to quantify the ratio of reaction rate from reaction kinetics to mass transfer rate. The Da number is the ratio of the reaction rate from reaction kinetics to mass transfer rate at initial reaction conditions. If the Da number less than 0.11, the reaction is the limit step. Figure 11 shows axial distribution of the Da number in RRZ, which indicates that the mass transfer rate between the core and annulus region is far less than the reaction rate from reaction kinetics, and the mass transfer between the core and annulus region becomes the limit step for the pyridine synthesis reaction. For such a small mass transfer rate, the corresponding conversion is low in RRZ. The riser diameter is 84 mm, which has more severe wall effect when compared with commercial risers. The severe wall effect leads to large annulus region along the riser radial direction, and in turn limits the mass transfer between the core region and annulus region, which at last results in low conversion in the RRZ. As for selectivity, as discussed in the Experimental Result section. The back-mixing in the riser is far less than that in the commercial reactors, which increase the selectivity of the products.

6 CONCLUSIONS A coupled fluidized bed reactor is proposed for gas phase pyridine synthesis. A pilot scale experimental apparatus was commissioned, and the product yield under different operating conditions was measured. A model is proposed to simulate the reaction process in the coupled reactor. During 15 days of continuous operating, there is no coke deposit on the feeding device, which indicates this novel feeding scheme is effective. The product yield in the coupled reactor reaches as high as that in the commercial reactors, which means enough gas-solid contact time is ensured by this coupled reactor. The selectivity of pyridine over 3-picoline is around 2.5 and is higher than that in commercial reactors, which is around 2.2. The model accurately predicts the product yield with an average deviation from experimental data of 12%. Based on model prediction, the conversion is mainly contributed by FZ and FRZ of the coupled reactor. 12


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Limited by the mass transfer and solid concentration in the riser, the RRZ contributes less than 10 percent to the total conversion.

SUPPORTING INFORMATION Tables with the production yield under different operating conditions and the value of reaction kinetic constants. ACKNOWLEDGMENTS This work was supported by NKBRDP (the National Key Basic Research Development Program) of China (973, 2012CB215000).

13



Ar a cA cD cE cF D Dg Fr Gs g h Hi Ho L k1 k2 k3 k4 kac kbd Lb m M N Nho P Pe R r rA rD rE s T t ug ups V v y z

NOMENCLATURE Archimedes number activity of catalysts, for fresh catalysts the value is unity concentration of acetaldehydes in feedstock (mol/L) pyridine concentration (mol/L) 3-picoline concentration (mol/L) coke concentration (mol/L) diameter (m) diffusion coefficient (cm2/s) Froud number catalysts circulation rate based on reactor cross sectional area (kg/m2/s) gravity constant, 9.8 N/kg height (m) the enthalpy of the feed stocks in the inlet of the computation node, (kJ/kg) the enthalpy of the feed stocks in the outlet of the computation node, (kJ/kg) the hear latent of the catalysts, (kJ/kg/K) reaction rate constant for reaction A to D (L/(mol.s)) reaction rate constant for reaction A to E (L/(mol.s)) reaction rate constant for reaction D to F (s-1) reaction rate constant for reaction E to F (s-1) mass transfer coefficient from core to annulus region (m/s) mass transfer coefficient from bubble phase to dense phase (m/s) length of bubbling bed reaction zone (m) mass flow rate (kg/s) mole mass (kg/kmol) number of CSTRs number of holes per cm2 in gas distributor pressure (Pa) Peclet number ideal gas constant 8.314 (J/(mol.K)) reaction rate (mol/(L.s) reaction rate for acetaldehydes (mol/(L.s)) reaction rate for pyridine (mol/(L.s)) reaction rate for 3-picoline (mol/(L.s)) product selectivity reaction temperature (K) residence time (s) superficial gas velocity (m/s) particle velocity, Gs/ρp (m/s) volumetric flow rate (m3/h) diffusion volume (cm3) pyridine and 3-picoline production yield height of the reaction zone (m)

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Greek symbols ηaf acetaldehydes to formaldehydes ratio ηam ammonia to aldehydes ratio △ difference σ standard deviation Фc the ratio of core area to cross area ρs solid phase concentration (kg/m3) ε solid fraction μg gas phase diffusivity (Pa.s) Subscripts a annulus ax axial c core ca catalyst d dense phase f fast fluidized bed fe feedstock g gas phase i in o out 0 initial pb packed bed pr product s solid phase tur turbulent fluidization

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Figures

1: Self adjusting flow measuring system Figure 1. Process of experiment apparatus

19



Figure 2. Axial positions of thermocouples and pressure gauges

Figure 3. Computation node for catalysts circulation rate

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A: Acetaldehyde; B: Formaldehyde; C: Ammonia; D: Pyridine; E: 3-picoline; F: Coke Figure 4. Model reaction scheme for pyridine synthesis

Figure 5. CSTR in series model

Experiment data 600

Calculated ρb (kg/m3)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


+15% 400

-15%

200

0 0

200

400

600

Experimental ρb (kg/m3) Figure 6. Comparison between the modified correlation prediction result and experimental data

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Figure 7. Core-annulus model

CA (mol/L)

0.008

CD (mol/L)

s: selectivity

CE (mol/L)

0.006

CF (mol/L)

Over reaction

s/1000 0.004

0.002

0.000

T=728 K, CA0=0.008 mol/L, ρs=200 kg/m3/s -1

0

1

2

3

4

5

t (s)

Figure 8. The selectivity variation with reactants residence time

22


6

7

8

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90

Experiment data 60

Calculated y

+15%

30

-15%

0 0

30

60

90

Experimental y

Figure 9. Comparison between the model prediction and experiment data Fluidized bed Reaction zone acetaldehyde pyridine 3-picoline coke

0.006

Riser Reaction Zone

0.004

c (mol/L)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


0.002

0.000

Feeding Zone

0

3

6

9

12

15

z (m)

Figure 10. Concentration variation along axial direction of the coupled reactor by model prediction

23



7

7

x 10

6

5

4

Da

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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3

2

1

0

1

1.5

2

2.5

3 z/0.1 (m)

3.5

4

4.5

5

Figure 11. The Da number distribution along axial direction in the RRZ

For Table of Contents Only

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Coupled fluidized bed reactor for pyridine synthesis - Industrial

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