Continuous Hydrogenolysis of N-Diphenylmethyl Groups in a Micropacked-Bed Reactor

Article abstract of DOI:10.1021/acs.oprd.9b00416

In recent years, with the advancements in continuous flow technology and the ever-increasing demand for green processes, continuous flow chemistry has become more and more widely adopted in the pharmaceutical industry. In this work, the continuous hydrogenolysis of N-diphenylmethylazetidin-3-ol to 3-azetidinol in micropacked bed reactors was demonstrated. The effects of different catalysts, solvent types, and the additives on the reaction in a micropacked-bed reactor were investigated. The results indicate that the reaction rate per reactor volume is increased by 100 times because of the larger interfacial area and shorter diffusion distance in micropacked reactors. To further study the long-term stability of the reaction system, the flow system was successfully operated for 240 h by adjusting the reaction temperature and liquid flow rate. The reaction kinetics model for the hydrogenation of N-diphenylmethylazetidin-3-ol in methanol was studied after the internal and external diffusion limitations were eliminated. In addition, the type of adsorption of the reactants on the catalyst and the rate-determining step of the reaction were investigated.

Full text of DOI:10.1021/acs.oprd.9b00416

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Full Paper
Continuous hydrogenolysis of N-diphenylmethyl
groups in a micro-packed bed reactor
Jiacheng Tu, Le Sang, Han Cheng, Ning Ai, and Jisong Zhang
Org. Process Res. Dev., Just Accepted Manuscript • DOI: 10.1021/acs.oprd.9b00416 • Publication Date (Web): 04 Dec 2019
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Continuous hydrogenolysis of N-diphenylmethyl groups in a
micro-packed bed reactor
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Jiacheng Tu ab, Le Sang a, Han Cheng a, Ning Ai b, Jisong Zhang a
a. State Key Laboratory of Chemical Engineering, Department of Chemical
Engineering, Tsinghua University, Beijing 100084, China.
b. School of Chemical Engineering, Zhejiang University of Technology, Hangzhou
310014, Zhejiang, China.
To whom correspondence should be addressed. Email: jiszhang@tsinghua.edu.cn
(J.S. Zhang)
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The Table of Contents
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Abstract
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In recent years, with the advancements in continuous flow technology and the ever-
increasing demand for green processes, continuous flow chemistry has become more
and more widely adopted in the pharmaceutical industry. In this work, the continuous
hydrogenolysis of N-diphenylmethylazetidin-3-ol to 3-azetidinol in micro-packed bed
reactors was demonstrated. The effects of different catalysts, solvent types and the
additives on the reaction in a micro-packed bed reactor were investigated. The results
indicate that the reaction rate per reactor volume in increased by 100 times due to the
larger interfacial area and shorter diffusion distance in micro-packed reactors. To
further study the long-term stability of the reaction system, the flow system was
successfully operated for 240 h by adjusting the reaction temperature and liquid
flowrate. The reaction kinetics model for the hydrogenation of N-
diphenylmethylazetidin-3-ol in methanol was studied after the internal and external
diffusion limitations were eliminated. In addition, the type of adsorption of the reactants
on the catalyst and the rate-determining step of the reaction were investigated.
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Keywords: Hydrogenation; Debenzylation; Kinetics; Micro-packed bed reactors
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Introduction
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The hydrogenation reactions are ubiquitous chemical transformations in
petrochemical industry, food industry and pharmaceutical industry. About 9% of all
chemical transformations in the pharmaceutical synthesis are hydrogenation reactions
1. In addition to the common hydrogenations of unsaturated bonds, deprotection
reactions with H2 (hydrogenolysis) are also widely used in organic synthesis,
accounting for about 3 % of total pharmaceutical synthesis 1.
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In the production of pharmaceuticals and fine chemicals, protection and deprotection
of functional groups such as alcohols and amines are often necessary to avoid possible
side reactions 2. In particular, formation and cleavage of benzyl-type ethers and benzyl-
3
type amines represent a common strategy in organic synthesis . The benzyl group is
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easily installed, by substitution of benzyl halide or condensation of benzaldehyde ,
and is stable to mild aqueous acid/base solutions, metal hydrides and mild oxidants 3a.
Later, it can be removed through direct hydrogenation with the gas of H2, transfer
hydrogenation with NaBH4 or using Lewis acids such as AlCl3 or triflic acid 6. Among
these methods, precious metal catalyzed hydrogenolysis with high pressure gas of H2
is one of the green and efficient options, in accordance with “atom economy”, “reduce
derivative” and “catalysis” described in the green chemistry principles 7.
The batch reactor is widely used for hydrogenation in the pharmaceutical industry
and fine chemical industry because the reactor structure is simple. However, there are
still several problems limiting its application. First, the heterogeneous hydrogenation is
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usually thought to be a fast reaction . Due to the severe gas-liquid mass transfer
limitation for hydrogenation in batch reactors, a long reaction time (3~24 h) and a large
reactor volume (1000~5000 L) are required in the industry production. Second, multiple
gas replacement operations and separation of the catalysts are needed for one batch
production, which is troublesome and labor-intensive. Third, safety issue is concerned
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for the hydrogenation process because of the store of a great amount pressurized
hydrogen and frequent operation of palladium carbon catalyst 9. As a result, it is highly
desired to develop a safe and efficient hydrogenation method for pharmaceutical
industry and fine chemical industry.
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The continuous flow technology has been identified as the number one research area
for the advancement of green pharmaceutics by the ACS GCI Pharmaceutical
Roundtable 10. Continuous flow processes have many inherent advantages including the
effective suppression of side reactions by precise control of reaction conditions,
reduced cost of the separation/purification process, increased production, reduced
energy consumption, reduced equipment volume, and excellent safety 11. Among them,
the continuous hydrogenation in micro-packed bed reactors (µPBRs) is promising for
the challenges of hydrogenation. In µPBRs, the pressurized hydrogen and liquid
reactants flow concurrently through packed beds of catalyst particles with diameters
usually smaller than 500 µm. Due to the microscale dimension, it has the advantages of
good gas-liquid-solid mass transfer, excellent heat transfer and improved safety while
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maintain the plug flow characteristics and fixed bed catalyst immobilization
.
Bimbisar Desai et al. achieved the selective debenzylation of substrate containing
acylamide and carbon-carbon double-bond structure in the H-Cube packed bed reactors
13. This reaction was performed with Pd/C catalysts under 1 bar, 40 ℃ and a low liquid
rate of 1.0 ml/min with the yield of 95%. Marie-Christine Matos et al. successfully
achieved simultaneous removal of multiple benzyl groups in µPBRs with the yield of
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89 % under 80 bar and 60 ℃ with Pd/C catalysts . These results demonstrate the
efficiency and safety of µPBRs in debenzylation than that of the batch reactor.
In this paper, a flow system based on µPBRs was developed to perform the H2-based
hydrogenolysis. The hydrogenolysis of N-diphenylmethylazetidin-3-ol (DMAOL) to 3-
azetidinol (Figure 1) was employed as the model reaction, which is an organic
pharmaceutical intermediate for the synthesis of antibiotics and antibacterial agents15.
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The catalysts and solvents suitable for this kind of reaction were first screened and then
the reaction performance and kinetics were investigated.
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Figure 1. Hydrogenolysis of N-diphenylmethylazetidin-3-ol to 3-zaetidinol
Material and methods
Chemicals. N-diphenylmethylazetidin-3-ol (C16H17NO) was acquired from
PharmaBlock Sciences Ltd. (Nanjing); Methanol (CH3OH, 99.9 %) and Acetonitrile
(C2H3N, 99.9 %) were purchased from J&K Scientific Ltd. (Beijing); Tetrahydrofuran
(C4H8O, AR), Ethyl acetate (C4H8O2, AR), Toluene (C7H8, AR) and Triethylamine
(C6H15N, AR) were purchased from Beijing Tongguang Fine Chemicals Company; 2-
Methyltetrahydrofuran (C5H10O, 99 %) was purchased from Energy Chemical;
Diphenylmethane (C13H12, 99 %) was purchased from Nuobeiwei Technology Co. Ltd.
(Beijing); Acetic acid (CH3COOH, 99.5 % ) was purchased from Beijing Chemical
Factory. Hydrogen (H2, 99.999%) was purchased from Beijing Beiwen Gas
Manufacturing Plant. Nitrogen (N2) was purchased from Air Liquide (Tianjin). The
Palladium hydroxide/Alumina particle catalyst (Pd(OH)2/Al2O3, 7 wt%) with an
average size of 500 µm and palladium/alumina particle catalyst (Pd/Al2O3, 5 wt%) with
an average size of 500 µm were acquired from Dalian Institute of Chemical Physics
The Palladium/Carbon powder catalyst (Pd/C, 5 wt%, 55 % water content) and
Palladium hydroxide/ Carbon powder catalyst (Pd/(OH)2/C, 5 wt%, 55 % water content)
were acquired from HEOWNS Biochem Technology LLC (Tianjin) and Macklin
Biochemical Co. Ltd. (Shanghai), respectively.
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Experimental setup. A schematic overview of the experimental setup is shown in
Figure 2. Liquid was delivered by a plunger pump (Beijing Oushisheng Technology
Co. Ltd.). N2 gas and H2 gas from the regulated cylinder were fed via a mass flow
controller (Beijing Sevenstar Electronics Co. Ltd.) The outlet of the pump was
connected with a safety valve (Beijing Xiongchuan Technology Co. Ltd.) to
automatically release the liquids from the tubing when the pressure exceeded 40 bar.
Two pressure gauges were placed in the inlet and outlet of the reactor to monitor the
pressure drop across the reactor. The reactor and two coils were immersed in a water
bath to control the reaction temperature. The outlet of the packed bed was connected to
a back pressure regulator (Beijing Xiongchuan Technology Co. Ltd.) to control the
system pressure.
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Figure 2. The schematic overview of the flow system for hydrogenation based on
the µPBRs.
The mixing of the feeding gas and liquid before the entering the packed bed reactor
greatly affects the stability and reproducibility of the system. To minimize the
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hydrodynamic instabilities, the reactor structure followed the design which was
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described in literature and more details about the reactor structure can be found in
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this paper. The packed tube is made of stainless steel with a length of 20 cm, an outer
diameter of 6.35 mm (1/4 inch) and an inner diameter of 4.35 mm.
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The reactant DMAOL was prepared as a specific concentration in the solvent. The
particulate catalysts were filled into the reactor through a funnel and the catalyst bed
was packed as tight as possible. The micro-packed bed packed with a certain mass of
catalyst was pre-wetted by the solution, then H2 was transported into the reactor at a
specific flow rate (F ). Back pressure (P) and liquid flow rate (
were then set to the
퐹퐿)
G
required values while simultaneously raising system temperature (T). After waiting for
the reaction time, which was at least three times of the liquid residence time ( ), the
휏
system achieved steady state and a sample was collected. When the desired samples
were acquired, water bath was turned off and system pressure was reduced to
atmospheric pressure. After the system was depressurized, N2 gas was transported
through the micro-packed bed reactor to fully remove the residual gas-liquid mixture
and ensure the catalyst was stored in a nitrogen atmosphere.
As hydrogenation in µPBRs involves gas-liquid-solid process, the residence time of
liquid
representing the reaction time should be carefully calculated. According to
휏
the definition of liquid holdup in our previous paper 16, is calculated by the following
휏
equation:
휋퐷2퐿휀ℎ
(1)
휏 =
퐹퐿
where, D (cm) and L (cm) are the inner diameter and the length of the packed tubing,
respectively; is the bed porosity; h is the liquid holdup representing the fraction of
휀
interstitial open space filled by liquid. From this equation, we can find that the residence
time is only determined by the liquid holdup, which can be calculated with the equations
reported in literature 16.
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Sample analysis. The samples were collected directly from the outlet of the
hydrogenation system. The samples were diluted and measured by HPLC (Agilent 1260
Infinity II) with a UV detector under the following conditions: the column temperature,
30 C; the injection volume, 5 μL; the ultraviolet wavelength, 220 nm; the mobile phase
composition, Acetonitrile: water (70:30, v/v), 1.0 mL/min; the column, Agilent Eclipse
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plus C18 (4.6
250mm, 5 ). No side reaction was observed when performing
μm
×
hydrogenolysis of DMAOL in the µPBRs. Hence the selectivity of the reaction is
assumed to be 100 %. The product of 3-zaetidinol cannot be measured by HPLC
because it has no UV absorption. Only the concentrations of DMAOL and
diphenylmethane are measured to calculate the conversion (X) of DMAOL as follows.
퐶퐵
푋 = 퐶 + 퐶 × 100%
(2)
퐴
퐵
퐶퐴
퐶퐵
(mol/L) are the concentration of DMAOL and
where,
(mol/L) and
diphenylmethane in the sample, respectively.
Results and discussion
Screening reaction conditions. As we know that the catalyst and solvent have a
great influence on the hydrogenation. To develop a most efficient continuous
hydrogenolysis in µPBRs, the screening of catalyst and solvent was first performed.
Effect of catalyst. It is well known that catalytic activity can be influenced by the
catalyst type and support. It has been reported that Pd on activated carbon is more
effective for debenzylation than Pd on oxide supports such as SiO2, Al2O3, TiO2 17. In
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addition, Pearlman's catalyst—Pd(OH)2/C
was commonly used in debenzylation
reaction, which demonstrated the best catalytic activity 19. In this study, the catalysts of
Pd and Pd(OH)2 on the supports of activated carbon and alumina were tested under the
same conditions. The reaction performances and the catalytic activities using different
catalysts are given in Table 1. Here the mass of the substrate converted per unit time
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and unit mass catalyst [g(raw material)*g(catalyst)-1*min-1] was used to evaluate the
catalyst activity.
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Table 1．The comparison of the catalytic activities of several catalysts on the
hydrogenolysis of DMAOL
Catalyst
Conversion
Catalyst activity [g(raw
material)*g(catalyst)-1*min-1]
0.31
Pd/C (5wt. %, 55% H2O
content)
36.9%
Pd/Al2O3 (5wt. %)
Pd(OH)2/C (5wt. %, 55%
H2O content)
28.3%
86.3%
0.11
0.73
Pd(OH)2/Al2O3 (7wt. %)
86.4%
0.23
Experimental conditions: flow rate of DMAOL (3.8 wt.% in methanol) solution: 0.3
ml/min; P= 21 bar; T= 313.15 K; FG= 40 sccm (the molar ratio of H2 to DMAOL: 43);
the amount of catalyst (diluted with 4.5 g alumina sphere): 0.5 g.
From Table 1, it can be concluded that the catalytic activity of Pd(OH)2 catalyst is
about two times of that of Pd catalyst for the hydrogenolysis reaction. At the same time,
it also proves that Pd or Pd(OH)2 catalyst on activated carbon is more efficient than Pd
or Pd(OH)2 on oxide supports of Al2O3. But the carbon support used here was at powder
state, which would cause a large pressure drop and even clogging in µPBRs. For the
oxide supports of Al2O3, no significant pressure drop was observed, which means that
values of the pressure transducers before and after the packed bed reactor were almost
the same. As a result, Pd(OH)2/Al2O3 catalyst was adopted for the following research
in this paper.
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Effect of solvent. In hydrogenation, the solvent frequently plays an important role
20, which would affect the solubility of hydrogen, adsorption mechanism of solvent
molecules on the active sites of the catalyst, agglomeration of catalyst particles and
nonbonding interactions between reactant or product molecules with the solvent 21. The
choice of a suitable solvent can effectively increase the reaction rate, and conversely, it
will inhibit the reaction rate and even poison the catalyst. Here, the effect of some
common organic solvents was explored. To ensure the catalyst has the same catalytic
activity, the reaction using the solvent of methanol was performed for a catalyst activity
test every time before the solvents were changed.
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Figure 3. Conversion of DMAOL in different solvents
Experimental conditions: flow rate of DMAOL (3.8 wt.% in methanol) solution: 0.5
ml/min; P= 21 bar; T= 313.15 K ; FG = 40 sccm (the molar ratio of H2 to DMAOL:
26); the amount of 7 wt. % Pd(OH)2/Al2O3 catalyst (diluted with 4.0 g alumina sphere):
1.0 g.
Figure 3 shows the effect of several organic solvents on the conversion of DMAOL
at the same experimental conditions. When the solvents are methanol and ethyl acetate,
X can reach about 81% and 85% at the residence time of 2.1 min in the micro-packed
bed, indicating the beneficial effect on the catalytic activity. Compared with the two
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solvents, the reaction rates in the other three solvents are much slower. In addition, it is
found that the X is abruptly decreased after the use of the tetrahydrofuran (Entry 2) and
2-methyltetrahydrofuran (Entry 8). It indicates that tetrahydrofuran substances have a
certain toxic effect on the catalyst which should be avoided in the reaction. But the
reason for this toxic effect is still unclear and we didn’t find any reports about the toxic
effect of tetrahydrofuran for the hydrogenolysis with Pd(OH)2 catalysts. Further
research are required to find the exact mechanism of catalyst deactivation. As a result,
both methanol and ethyl acetate are excellent solvents for hydrogenation of DMAOL.
Considering the low price and easy recyclability of methanol (low boiling point),
methanol is chosen as the solvent for this reaction.
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Effect of additive. It is known that N-benzyl is the acid-labile protecting group.
Acidic conditions favor N-debenzylation, while basic conditions inhibit N-
debenzylation 22. In order to investigate whether the effect of acidic or basic additives
on the removal of the diphenylmethyl group was similar to that of the debenzylation,
an experiment was carried out by adding acetic acid and triethylamine to the solution.
The conversions with different additives are given in Table 2.
Table 2．The effect of additive on conversion of DMAOL
Entry
Pressure/bar
Additive
No additive
Acetic acid
Triethylamine
Water
Conversion/%
79.7
1
2
3
4
21
21
21
21
91.9
67.6
73.5
Experimental conditions: flow rate of DMAOL (3.8 wt.% in methanol) solution: 0.5
ml/min; T= 313.15 K ; FG= 40 sccm (the molar ratio of H2 to DMAOL: 26); the amount
of 7 wt. % Pd(OH)2/Al2O3 catalyst (diluted with 4.0 g alumina sphere): 1.0 g; the
concentration of additive is 8.8 wt. % in entry 2, 3 and 4.
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It is found that the conversion of entry 2 is higher than that of entry 1 while the
conversion of entry 3 is lower than that of entry 1. The results indicate that acidic
conditions favor the removal of the diphenylmethyl group, while basic conditions
inhibit the removal of the diphenylmethyl group. The results indicate that the
debenzylation with hydrogenolysis can be enhanced by addition of acetic acid into the
solution and the addition of basic additive should be avoided. The conversions of entry
4 is also slightly lower, indicating that the existence of water would have an adverse
effect on the reaction. The exact reason for the adverse effect of water on the reaction
is still unclear, which required further investigation.
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Reaction performance test. In this part, we compared the reaction performance of
traditional batch reactor (250 mL volume) with that of micro-packed bed reactor. Here
the mass of the substrate converted per unit time of the unit reactor volume [g(raw
material)*mL(reactor volume)-1*min-1] was used to evaluate the reactor performance.
As shown in Table 3, the reaction rate in batch reactor was very slow with a low
conversion of 18.5% after 3h and 38.1% after 5h. The main reason is the limited gas-
liquid mass transfer rate in batch reactor. In contrast, due to the larger interfacial area
and shorter diffusion distance in micro-packed reactors, a high conversion of 63% can
be obtained with only a short residence time of 1.9 min. In addition, the value of
conversion per unit volume per unit time of the micro-packed bed reactor is 100 times
of that of the batch. As a result, the micro-packed bed reactor is more efficient for the
hydrogenation of DMAOL than the traditional batch reactor.
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Table 3．The comparison of performances of traditional batch and micro-packed bed
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Entry
Residence Conversion/% reactor performance [g(raw material)*
time
3.0 h
mL(reactor volume)-1*min-1]
2.5 × 10-5
1
2
3
18.5
38.1
63.0
5.0 h
3.1 × 10-5
1.9 min
3.1 × 10-3
Experimental conditions: initial concentration of DMAOL in methanol: 3.8 wt.%;
T= 333.15 K; P= 21 bar; the amount of Pd(OH)2/Al2O3 catalyst (7 wt. %): 0.5 g; the
entry 1 and 2 were carried out in traditional batch: stirring speed was 200 rpm to avoid
the damage of catalysts; raw material volume was 200 mL; the entry 3 was carried out
in micro-packed bed reactor: FL = 0.5mL/min, FG= 40 sccm (the molar ratio of H2 to
DMAOL: 26).
To further study the long-term stability of the reaction system and catalyst life of
Pd(OH)2/Al2O3, the reaction performance was investigated in the micro-packed bed
reactor for 240 h as shown in Figure 4. In order to achieve the maximum reaction
conversion, the catalyst bed was not diluted with alumina sphere filler (4.3 g
Pd(OH)2/Al2O3) and the concentration of the substrate was further improved to 10 wt.%
in this experiment. As first, the fresh catalyst had a high catalytic activity and the
conversion could reach 100% for 70 h. After 70 h, the conversion began to decrease
because of the deactivation of the catalyst. The reason of deactivation and following
regeneration method are still unclear and we are still working on them. The most
possible reason for the deactivation is the adsorption of substances on the precise metal
catalyst due to the strong interaction between amine and precise metal. To increase the
reaction rate, the reaction temperature was gradually increased to maintain the high
conversion. After 116 h, the liquid flow rate was decreased from 0.4 ml/min to 0.3
m/min to increase the residence time to ensure the conversion. By increasing the
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temperature and decreasing the flow rate, the conversion of the substrate could stay
above 99.5% for 190 h. After that, the conversion decreased rapidly indicating a severe
catalyst deactivation occurred and the catalyst was required to be replaced or
regenerated. From the results in Fig. 4, it demonstrates that about 90 g DMAOL can be
efficiently converted to the product for one gram Pd(OH)2/Al2O3 catalyst in the micro-
packed bed and the continuous flow system can achieve an efficient production of 3-
azetidinol for 190 h (the temperature and liquid flowrate were changed after 70h).
However, more issues such as heat transfer, gas-liquid mixing and pressure drop across
the packed bed should be carefully considered in the scaling up of this continuous
system.
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Figure 4. The stability of reaction system in the micro-packed bed reactor
Experimental conditions: initial concentration of 1-(diphenylmethyl)-3-
hydroxyazetidine: 10 wt.%; P= 25 bar ; the amount of Pd(OH)2/Al2O3 catalyst (7 wt. %):
4.3 g. When the reaction time is less than 116 h, FL= 0.4 mL/min, FG= 60 sccm; when
the reaction time is more than 116 h, FL= 0.3 mL/min, FG = 45 sccm.
Kinetic study of hydrogenolysis of DMAOL
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Mass transfer considerations. In the three-phase catalytic reaction system, the mass
transfer typically includes three steps in series: dissolution of H2 in liquid, diffusion of
dissolved H2 through liquid to the solid catalyst surface and intra-particle diffusion to
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the active catalytic sites, which can influence the reaction kinetics . Therefore, it is
important to confirm that the mass transfer limitation can be ignored before determining
the kinetic parameters.
First, external mass transfer should be carefully evaluated 10a. Due to the use of pure
hydrogen as reactant, the mass transfer resistance occurred in the gas film side of the
gas–liquid interface can be ignored, and the mass transfer resistance on the liquid side
at gas–liquid interface and the mass transfer resistance at the liquid–solid interface
should be considered24. Exploring the effect of different liquid superficial velocities
(flow rate/cross-sectional area of packed bed) on the reaction rate at the same residence
time can be used to determine whether external mass transfer limitation can be ignored.
In the present study, four µPBRs with the same volume (3.0 ml) and different internal
diameters (4.35 mm; 4.57 mm; 5.35mm; 5.75mm) were designed for the experiments.
These reactors were filled with the same amount of catalyst. This method can
successfully achieve different liquid superficial velocities in different reactors under
the same flow rate and residence time.
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Figure 5. Effect of different liquid superficial velocities on the average reaction rate.
Experimental conditions: initial concentration of DMAOL: 5.7 wt. %; P= 21 bar; T=
313.15 K; FG= 40 sccm; the amount of 7 wt. % Pd(OH)2/Al2O3 catalyst (diluted with
4.0 g alumina sphere): 1.0 g.
Figure 5 shows that when the liquid superficial velocity exceeds 5.0 cm/min, the
average reaction rate does not change significantly as the liquid velocity increases. This
result indicates that the reaction is independent of external mass transfer when the liquid
superficial velocity is greater than 5.0 cm/min. It can be considered that the external
mass transfer resistance has been eliminated.
Internal mass transfer limitations within the porous media were evaluated by
calculating the Weisz modulus (Mw), which ratios the effective reaction time and the
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diffusion time within the catalyst particle . The Weisz modulus of substrate and H2
were calculated as shown in the supporting information. The values of the Weisz
modulus were determined to be 0.123 and 0.776, for H2 and DMAOL, respectively,
indicating only minor contributions from intraparticle diffusion and the internal mass
transfer resistance can be ignored.
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Kinetic model of hydrogenolysis of DMAOL. The reaction kinetic model was
studied after the internal and external transfer resistances were eliminated to ensure that
the reaction was in the kinetic-control region. Here, the kinetics is presented by a power-
law model, which can be described by an empirical kinetic expression:
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r = 푘′퐶훼퐴푃훽 = ― 푑퐶퐴
(3)
푑푡
where
are the reaction orders of reactants.
훼 and 훽
First, the value of
was determined by changing the initial concentration of the
α
reactants at the same residence time. During the experiment, it was found that there was
no significant pressure drop in the reactor through the pressure gauges at the inlet and
outlet of the packed bed reactor, the reaction pressure can be considered to be constant.
푘′푃훽
At first, the temperature and pressure were kept unchanged, the term
could be
regarded as a constant K, and the following equation is obtained:
푟 = 퐾퐶훼퐴 = ― 푑퐶퐴
(4)
푑푡
After integrating the Equation 4 and considering the boundary conditions [t= 0, CA=
C ; t= , C = C (1-X)], Equation 5 is obtained.
휏
A0
A
A0
1
푋 = 1 ― 퐾휏(훼 ― 1) + 1
1 ― 훼
(5)
[
]
1 ― 훼
퐶퐴0
Figure 6. The conversion versus the initial concentration of DMAOL in methanol.
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Experimental conditions: P= 21 bar; T= 313.15 K; FG= 40 sccm; FL = 0.9 mL/min;
5
6
63.6 s; the amount of 7 wt. % Pd(OH)2/Al2O3 catalyst (diluted with 4.0 g alumina
τ =
7
8
sphere): 1.0 g.
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The hydrogenolysis of DMAOL was performed at nine initial concentrations ranging
from 0.033 mol/L to 0.33 mol/L and the conversions at different concentrations were
given in Figure 6. The experimental data was fitted with Equation 5. It was found that
when α is 0.53, the calculated curve agreed well with the experimental values (R2=0.99),
indicating that the reaction order of DMAOL is 0.53.
The value of
is determined by changing the reaction pressure at the same
훽
residence time. The liquid holdup is related to the superficial velocities of the gas and
liquid. The change of the reaction pressure will lead to an obvious change for the
superficial velocity of the gas and corresponding residence time of liquid. To eliminate
the effect of reaction pressure on the residence time, the liquid flowrate was varied from
0.58 to 1.04 mL/min according to the model described in literature 16 to ensure the same
residence time. Taking α
0.53, integrating and deforming the Equation 3, the
=
Equation 6 was obtained.
1
훽
′
0.47
푋 = 1 ― 1 ― 0.47푘 휏푃
(6)
[
0.47
]
퐶퐴0
Figure 7. The conversion versus the reaction pressure.
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Experimental conditions: CA0= 0.20 mol/L; T= 313.15 K; FG= 40 sccm; flow
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6
velocity = 16.32 cm/min;
73.5 s; the amount of 7 wt. % Pd(OH)2/Al2O3 catalyst
τ =
7
8
(diluted with 4.0 g alumina sphere): 1.0 g.
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The hydrogenolysis of DMAOL was performed at nine reaction pressures and the
obtained experimental results were shown in Figure 7. Then the experimental data was
fitted with Equation 6. It shows that when
is 0.31, the calculated curve agreed well
훽
with the experimental values (R2= 0.97), indicating that the reaction order of hydrogen
pressure is 0.31.
After determining the values of
and , the two values were substituted into
훼
훽
′
푘
Equation 3, Equation 7 can be obtained to calculate the value of
.
0.47
0.47
[
]
1 ― (1 ― 푋)
푘′ = 퐶
퐴0
(7)
0.47푃0.31
휏
(a)
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(b)
′
Figure 8. (a) The conversion versus the reaction temperatures. (b) Plot of ln versus
k
1/RT.
Experimental conditions: CA0= 0.20 mol/L; P=21 bar; FG = 40 sccm; FL = 0.9mL/min;
63.6 s; the amount of Pd(OH)2/Al2O3 catalyst (7 wt. %): 1.0 g.
휏 =
Figure 8. (a) shows the conversions at varied reaction temperatures ranging from
303.15 K to 343.15 K. The results indicate that the reaction rate increases as the
푘′
temperature increases. The values of
at different temperatures were calculated with
― 퐸
푎
푅푇 , the activation energies
푘′ = 퐴푒
the Equation 7. According to the Arrhenius formula,
′
푘
(Ea) and Arrhenius constant (A) can be obtained by fitting ln and 1/RT as shown in
Figure 8. (b), which are 18.6 ± 0.7 kJ/mol and 2.68 ± 0.25 mol0.47/(L0.47bar0.31s),
respectively (R2=0.99).
After all the parameters were obtained, the kinetic model could be established.
4
―1.86 × 10 (퐽/푚표푙)
푚표푙0.47
퐿0.47 ∙ 푏푎푟 ∙ 0.31푠 ∙ 푔cat
―1
―1
(
퐽 ∙ 푚표푙
∙ 퐾
)
× 푒8.314
× 푇퐶0퐴.53푃0.31
(8)
푟 = 2.68
(
)
The type of adsorption of the reactant on the catalyst and the rate-determining step
of the reaction were studied after the kinetic model was established. The heterogeneous
hydrogenation rate is determined by one of the steps of adsorption of the substrate or
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hydrogen, desorption of the product, and surface reaction when the reaction is in the
kinetics-control zone 26. In addition, due to the different adsorption mechanisms of the
reactants on the porous catalysts, research work on kinetic models has been further
complicated. Based on Langmuir equation, many rate expressions were studied in this
paper.
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Due to the different adsorption mechanisms of the reactants , there are up to 16
different theoretical kinetic models. All the possible theoretical models have been fitted
with the experimental data. If a mathematical model agrees well with the data, the
model can be considered to be the reaction kinetic model.
After excluding the other 15 models, the kinetic model was obtained (R2=0.99 for
CA0, R2=0.96 for P, the detailed process and fitting results were shown in supporting
information):
푘퐾퐴퐶퐴퐾퐻푃
(9)
푟 = (
2
퐾 푃
)
(
)
1 + 퐾퐴퐶퐴 + 퐾퐵퐶퐵 1 +
퐻
where, k, KA, KB, and KH are the reaction rate constant, adsorption equilibrium constant
of substrate, product and H2, respectively.
From the kinetic model, we can find that the diphenylmethyl removal rate of
DMAOL with Pd(OH)2/Al2O3 catalyst was controlled by surface reaction rate and the
adsorption type of substrate and H2 on the catalyst are non-competitive and dissociative
22b
adsorption. Hyun Tae Hwang
et al. studied the debenzylation of N-benzyl-4-
fluoroaniline, and the obtained rate-determining step and the adsorption type of the
reactant on the catalyst were consistent with the conclusions in our study. In addition,
compared with the activation energy value (138.4 kJ/mol) obtained by Hyun Tae
Hwang, we found that the activation energy value (18.6 kJ/mol) in our study is very
small, which indicate that the palladium hydroxide catalyst have better catalytic activity
for debenzylation than palladium catalyst. The kinetic model obtained in this study
maybe helpful for the reactor design and process optimization for the continuous
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deprotection development.
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Conclusion
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In this study, a micro-packed bed reactor has been successfully applied to the
continuous hydrogenation of DMAOL using hydrogen gas. It is demonstrated that the
µPBRs is more efficient for the hydrogenation of DMAOL than the traditional batch
reactor. The conversion per unit reactor volume per unit time in the micro-packed bed
reactor is 100 times of that of the batch. The reaction kinetics model was obtained after
the external and internal mass transfer limitations have been eliminated and the reaction
is in the kinetics-control region. The reaction rate is found to be 0.53 order with respect
to DMAOL concentration and 0.31 order to hydrogen pressure. The activation energy
(Ea) of the rate constant is 18.6 kJ/mol and the Arrhenius constant (A) is 2.68
mol0.47/(L0.47bar0.31s). The experimental results were fitted with several kinetic models
to determine the rate-determining step, which indicated that the diphenylmethyl
removal rate of DMAOL with Pd(OH)2/Al2O3 catalyst was controlled by surface
reaction rate and the adsorption type of substrate and H2 on the catalyst are non-
competitive and dissociative adsorption.
Acknowledgements
We gratefully acknowledge the supports of the National Natural Science Foundation of
China (21978146), Tsinghua University Initiative Scientific Research Program
(2019Z08QCX02) and the State Key Laboratory of Chemical Engineering (SKL-ChE-
17T01) on this work.
Nomenclature
FL
FG
Liquid flow rate
Gas flow rate
ml/min
sccm
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P
T
Pressure
bar
Temperature
K
7
8
h
Liquid holdup
−
9
D
Reactor inner diameter
Reactor length
cm
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L
cm
CA
CB
CA0
X
Concentration of DMAOL
Concentration of diphenylmethane
Initial concentration of DMAOL
conversion of N-diphenylmethylazetidin-3-ol
Activation energy
mol/L
mol/L
mol/L
m/s2
kJ/mol
Ea
A
푚표푙0.47
Arrhenius constant
퐿0.47 ∙ 푏푎푟 ∙ 0.31푠 ∙ 푔cat
Greek letters
휏
Liquid residence time
Porosity
s
휀
−
Subscripts
G
L
Gas
Liquid
Supporting Information
The calculation of Weisz modulus and detailed results on the intrinsic kinetic model
fitting are supplied as Supporting Information.
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