Design and Comparison of Tubular and Pipes-in-Series Continuous

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Design and Comparison of Tubular and Pipes-in-Series Continuous Reactors for Direct Asymmetric Reductive Amination Martin David Johnson, Scott A May, Brian D. Haeberle, Gordon Randy Lambertus, Shon R. Pulley, and James R Stout Org. Process Res. Dev., Just Accepted Manuscript • DOI: 10.1021/acs.oprd.6b00137 • Publication Date (Web): 14 Jun 2016 Downloaded from http://pubs.acs.org on June 30, 2016

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Organic Process Research & Development

Design and Comparison of Tubular and Pipes-in-Series Continuous Reactors for Direct Asymmetric Reductive Amination Martin D. Johnson,†,* Scott A. May,† Brian Haeberle,

††

Gordon R. Lambertus,† Shon R.

Pulley,††† James R. Stout# †

Eli Lilly and Company, Indianapolis, Indiana 46285, United States

††

Synergy Industrial Corporation, Brookfield, WI, United States

†††

#

Elanco, Eli Lilly ＆ Co. Greenfield, United States

D&M Continuous Solutions, LLC, Greenwood, IN 46143, United States

ACS Paragon Plus Environment


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TOC Graphic


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Abstract A high pressure, high temperature Direct Asymmetric Reductive Amination (DARA) reaction was run in three different continuous reactor types; coiled tubes, horizontal pipes in series, and vertical pipes in series. The goal of this work was to design a continuous reactor scalable to large volumes needed for commercial scale manufacturing. With about 4000:1 substrate to dissolved catalyst ratio, the DARA reaction achieved about 90% conversion to desired product in 12 hours, about 8% conversion to undesired ketone, and about 95% ee, in batch and in all 3 types of continuous reactors. The 3 continuous reactors were compared on the basis of vapor liquid mass transfer rates, overall pressure drop, % liquid filled, and axial dispersion. Vertical pipes in series were selected for scale up to manufacturing because of scalability and vapor liquid mass transfer rates, although horizontal pipes had lower pressure drop and less potential for surging. Key words continuous reaction, plug flow reactor, pipes in series, direct asymmetric reductive amination, high pressure hydrogen Introduction In our previous work, we communicated the design, development and scale up of coiled tube reactors for high pressure asymmetric hydrogenation.1 This simple, low cost reactor type worked extremely well for homogeneously catalyzed continuous hydrogenations. However, the main limitation was scalability. The largest coiled tube continuous reactor that we have constructed and operated at Eli Lilly and Company is 73L, and our judgment is that the largest practical volume for this reactor design is about 200 L because of difficulty to bend fabricate, and support the larger diameter tubing, and the need for a custom agitated vessel for temperature



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control. The goal of this work was to design and develop an alternative reactor scalable to 2000 L. It would be used for continuous reactions with 12 hour mean residence time (τ) with volumetric throughputs up to 4000 L/day. Straight pipes in series connected by small diameter tubing jumpers was envisaged as a more scalable concept. A reactor of this design could easily reach desired volumes and could be contained in a compact heating shell similar to a shell in tube heat exchanger. Vertical bubble flow pipes in series and horizontal bubble flow pipes in series were both proposed. It was hypothesized that (i.) the straight pipes in series would be scalable to larger volumes than the coiled tubes, (ii.) the straight pipes would be more cleanable and inspectable than the coiled tubes, (iii.) horizontal pipes and coiled tubes would have lower overall reactor pressure drop than vertical pipes, (vi.) vertical pipes would have higher vapor liquid mass transfer rates than horizontal pipes or coiled tubes, and (v.) reaction solution would re-saturate with gas in the connecting tubing between pipes. In this work, all 3 reactor types (Figure 1) were used for continuous direct asymmetric reductive amination (DARA) reactions at research scale (68 ml to 780 ml) and the pipes reactors at pilot scale (11 L to 32 L). We also documented work on continuous hydroformylation in research scale pipes in series, with a total liquid volume in the reactor about 10 ml. 2 It was a scale down version of a 360 L vertical bubble flow pipes in series reactor in Lilly GMP manufacturing. 3 This continuous DARA paper is the documents the work that fills the gap between the previous papers on coiled tubes and pipes in series. This DARA chemistry was the project on which our reactor designs evolved from coiled tubes, to horizontal pipes in series, and finally to vertical pipes in series. This is the work that quantified each type of continuous reactor and compared the designs on the basis of scalability, pressure drop, vapor liquid mass transfer rates, and axial dispersion.


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Figure 1. Three Continuous Reactor Types Used by Eli Lilly for V/L Reactions

The flow reactor should designed to enable use of low catalyst loadings such that reaction rate is limited by catalyst loading, thus the need for 12 hour τ.4 Residence time distribution (RTD) for the continuous reactor should be narrow by designing for low axial dispersion number (D/uL). D/uL should be less than 0.01 so that impact of axial dispersion on conversion versus τ is minimal, such thatτ more closely matches batch reaction times and so that reactor volume required for full conversion is minimized. Lower axial dispersion also minimizes the amount of transitional material during startups and shutdowns. Vapor/liquid mass transfer rates should be sufficiently high so that conversion versus time matches a well-mixed batch autoclave. A target kLa should be greater than 0.001 s-1 so that mass transfer rate is sufficiently higher than reaction rate, which is typically 0.0001 to 0.0005 s-1 for this type of reaction. kLa is overall vapor liquid mass transfer coefficient, which is used to quantify the rate of gas dissolving into the liquid and is described more later in the paper. The reactor should run at least 95% liquid filled to minimize volume of hazardous reagent gas in the reactor at any one time. Pressure limits for



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continuous reactor system should be at least 70 bar. Overall pressure drop should be less than 10 bar from inlet to outlet. The reactor internals should be cleanable and inspectable. Continuous reactors are advantageous for high pressure vapor/liquid reactions5 in the pharmaceutical industry because of safety, cost, and quality. There are many examples of reduced cost and improved safety and quality of continuous reactions compared to batch.6 Hydrogenation in particular is a high risk type of reaction, because if the H2 is released it can result in highly flammable or explosive mixtures with air, which can lead to a fire or explosion even without a known ignition source.7 Two ways to mitigate this risk is to reduce the amount of H2 gas in the reactor headspace by utilizing a continuous reactor that operates almost completely liquid filled, and to operate the sealed continuous reactor outside the manufacturing module. Continuous reactors made of pipes and tubing are relatively low cost for achieving high pressure rating especially when compared to the cost of a high pressure autoclave. For example, the 32 L pipes in series reactor used in this study, pressure rated to >70 bar, only cost $3000 to fabricate (not counting the pumps and control system), while a 100 L batch autoclave rated to 70 bar and capable of the same weekly throughput would cost orders of magnitude more. Continuous processing has made great strides in pharmaceutical development manufacturing, and academic research over the past decade.8 Quality control advantages of continuous processing can be realized because of steady state operation, and real-time product quality information by on-line process analytical technology (PAT). PAT can help identify trends over time which can be used for real time adjustments to keep product quality high at all times and also maximize yield. This aligns with quality by design (QbD) principles, and implementation of continuous processes is supported by the FDA9. Results and Discussion


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Scheme 1 Direct Asymmetric Reductive Amination (DARA) Reaction

The DARA reaction between dimethylketal 1 and aminotetrazole 2 to form amine product 3 is shown in Scheme 1 and was first developed in batch reaction experiments. In addition to the desired product 3, the ketone hydrolysis product (4), intermediate imine (5) and enamine (6), opposite enantiomer (7) and alcohol (9) were also tracked. Analysis of reaction profiles are somewhat complicated by competing hydrolysis of ketal 1, imine 5, enamine 6 and methyl enol ether 8 to ketone 4. Ketone 4 reacts very slowly in the DARA reaction and extended reaction times result in minimal additional product. In the following discussion on reaction performance in batch and flow, consumption of ketal must be balanced with the extent of hydrolysis observed. Reaction completion can be associated with the amount of ketal 1 remaining, since ketone is only sparingly reactive and the HPLC method results in minimal hydrolysis of 1, 5¸or 6. However, the amount of product 3 produced in the reaction is the best gauge of performance.

Reaction Performance in Batch Reactors



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The original conditions developed for this reaction involved toluene as the solvent. This posed a challenge for continuous operation, since the feed was not a solution at room temperature in toluene. One of the initial goals was to achieve homogeneous solution at room temperature for reagents and products. It is operationally simpler for a continuous reactor if the feed and product streams do not need to be heated. We explored alternative solvent ratios using toluene, THF, and 2-MeTHF. Reactions were each run for 16 hours in an Endeavor in an oxygen free glove box. The reaction was catalyzed by homogeneous iridium catalyst ([Ir(COD)Cl] 2) and (S)-XylBINAP ligand. Results are shown in Table 1. Conversion to ketone (4) was high because starting ketal contained ~ 9-10 area% ketone. Hydrogen update data for the Endeavor tubes showed that reaction time was about 5 to 10 hours for the reactions at 115 °C and about 2 hours at 130 °C. Product mixtures from 0.8 M reactions led to solids being formed upon standing. Table 1. DARA results from batch experiments to investigate solvent composition and concentration. a

Solvent Ratio

Conc. (M)

Temperature

3 (%)

4 (%)

1 (%)

5+6 (%)

THF

0.8

115

71.4

23.09

0.94

THF

0.4

115

71.04

22.71

0

Tol/THF 50:50

0.8

115

72.66

21.78

0

Tol

0.8

115

74.71

20.59

0

Tol/THF 70:30

0.4

115

77.63

18.55

0

Tol/THF 50:50

0.4

115

71.63

24.32

0

Tol/THF 70:30

0.8

115

70.78

24.27

0

THF/2-MeTHF 50:50

0.8

115

75.92

19.11

0

Tol

0.4

115

75.08

15.56

0

0.4

115

79.06

15.84

0

Tol/THF 70:30 Tol/THF 70:30

b

0.4

115

74.81

19.95

0

Tol/THF 80:20

0.4

115

77.31

15.55

0

Tol

0.4

115

67.28

14.03

Tol/THF 80:20

0.6

130

70.89

18.81

0

Tol/THF 70:30

0.4

130

78.98

13.54

0

0.4

130

71.85

12.19

0

Tol/THF 70:30

b

11.3

a

0

Reaction conditions were ketal 1 (1 eq), aminotetrazole 2 (1.1 eq), CSA (0.02 eq), TBAI (0.01 eq), 4000 S:C, H2 (400 psig).


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b

These 2 batch reactions used 0.02 equivalents TBAI and 0.03 equivalents CSA.

Batch reaction rate was investigated at pressures 500 and 1000 psig H2 by using a 50 ml batch autoclave and taking samples at intermediate time points. Results are shown in Figure 2. Reaction conditions were ketal 1 (1 eq), aminotetrazole 2 (1.1 eq), S:C 4000, 7 volumes THF solvent, 116 °C, CSA (0.02 eq), and TBAI (0.01 eq). Stir rate in the 50 ml PARR autoclave in the range 161 to 315 rpm had a small impact on rate of conversion to product. Final conversion to ketone (4) in these experiments was 7.5 to 8.3% after 24 hours, and final chirality measurements were 94.8 to 95.2% ee. Figure 2 shows reaction model fits to the data points using first order rate constants. Reaction rate was about 0.33 to 0.35 h-1 at 500 psig and 0.5 to 0.55 h-1 at 1000 psig.

100 90 80 70 area % product (3)

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60 50

1000 psig, 315 rpm stir rate k=0.55/h, reaction at 1000 psig, 315 rpm

40

1000 psig, 162 rpm stir rate 30

k=0.5/h, reaction at 1000 psig, 162 rpm 500 psig, 315 rpm stir rate

20

k=0.35/h, reaction at 500 psig, 315 rpm 500 psig, 161 rpm stir rate

10

k=0.33/h, reaction at 500 psig, 161 rpm 0 0

5

10

15

20

Time (hours) Figure 2. Batch reaction rate as a function of pressure and stir rate. 50 ml autoclave. a


25

30


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a

Reaction conditions were ketal 1 (1 eq), aminotetrazole 2 (1.1 eq), S:C 4000, 7 volumes THF solvent, 116 °C, CSA (0.02 eq), TBAI (0.01 eq).

Reaction Performance in Coiled Tube Reactors DARA reaction 68 mL coiled tube continuous reactor. The DARA reaction was run continuously in a 68 mL coiled tube reactor made from 316L stainless steel tubing. The reactor tube was 3.175 mm o.d. 1.75 mm i.d., and 28 m long. It was coiled and submerged in a constant temperature oil bath. A picture of the reactor and heating bath is shown in the Supporting Information. Results of the continuous reaction experiments are shown in Table 2. The data in the table represented 3 different continuous reaction experiments. Entries 1 and 2 were from the same experiment. It was encouraging that 8 hour τ and 4000 S:C (Entry 4) gave 93.9% desired product (3) and 95% ee in the flow tube reactor, with only 5.68% ketone. Compared to the batch rate data in Figure 2, rate of conversion to product was similar between the continuous reactor and the well-mixed batch autoclave, perhaps a little faster in the continuous reactor. It was 93.9% product in flow versus 84.5% product batch, both at 8 hour time point. This last set of conditions (entry 4) was run for 79 hours of continuous flow for the purpose of getting an isolated yield from the process. Reaction product solution from the 79 hour flow run was worked up and isolated as a batch. Isolated product (3) yield was 80%, and solids were 99.9% potency and 99% ee. This was an encouraging result for the continuous reactor. Table 2. Experimental results for DARA reaction 68 mL coiled tube continuous reactor. a Entry

Solvent

τ (h)

S:C

3 (%)

ee (%)

4 (%)

1

40/60 THF/toluene

6

2000

92.4

-

7.3

2000

93.1

-

6.53

2000

90.8

2 3

40/60 THF/toluene Pure THF

8 6

7.75

93.9

4 60/40 THF/toluene 8 4000 95.0 5.68 Reaction conditions were ketal 1 (1 eq), aminotetrazole 2 (1.1 eq), 8 volumes solvent, 115 °C, CSA (0.02 eq), TBAI (0.01 eq), H2 (500 psig), 1/1 gas/liquid volumetric flow ratio at reactor outlet before depressurizing. a


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DARA reaction 780 mL coiled tube continuous reactor. The reaction was next scaled up in a 780 mL coiled tube reactor. The 316L stainless steel reactor consisted of two concentric tubes in series that were each 6.35 mm o.d., 4.57 mm i.d., and 24 m long. These were joined by a connecting tube that was 3.175 mm o.d. 1.75 mm i.d., and 0.4 m long. A picture of the reactor is shown the Supporting Information. The coiled tube reactor was submerged in a constant temperature oil bath for temperature control at 115 °C. Hydrogen gas flow was controlled so that gas/liquid volumetric flow ratio was 0.25/1.0 near the exit of the reactor, just before the depressurization and vapor liquid separation. Thus excess H2 was lower in the 780 ml coiled tube than in the 68 ml coiled tube. A later section of this paper shows that this lowers vapor liquid mass transfer rate. Experimental results are shown in Table 3. One of the goals of this experiment was to reduce the catalyst loading, therefore it cannot be directly compared to the 68 ml coiled tube reactor experiments or the batch experiments. However, later in this paper a side by side comparison was made between the 780 ml coiled tube reactor and 780 ml pipes in series reactor, and that reactor was run at comparable conditions to the batch reaction experiments. The three entries in Table 3 were all from the same continuous reaction experiment. The experiment began with 10,000:1 S:C and 15 hour τ. Conversion to desired product (3) was only 88.3%, therefore the catalyst pump flow rate was increased so that catalyst loading was 8000:1 S:C. The higher catalyst loading increased area% product (3) to 89.8%. Finally, holding the reaction solution in the tube reactor for 48 hours resulted in higher conversion to 92.1 area% product. Thus, at 8000:1 S:C, τ longer than 15 hours is needed to reach highest possible product concentration. Table 3. Experimental results for DARA reaction 780 mL coiled tube continuous reactor. a



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entry

τ (h)

S:C

3 (%)

4 (%)

ee (%)

1

15

10,000

88.32

9.0

96.5

2

15

8000

89.8

7.9

96.5

3 48* 8000 92.1 7.9 96.6 Reaction conditions were ketal 1 (1 eq), aminotetrazole 2 (1.1 eq), 6.5 volumes THF, 115 °C, 1000 psig H2, CSA (0.02 eq), TBAI (0.01 eq), a

Residence Time Distribution in Coiled Tube Reactor Residence time distribution was quantified experimentally by sampling during startup transition (Figure 3) and shutdown transition (Figure 4) and analyzing concentration of product in solution. The reactor was initially filled with only solvent, therefore startup transition was from solvent to full concentration. The opposite was done during shutdown transition, where pumps were switched to solvent at the same flow rates and continued to pump until no product was detected in reactor effluent. In the figures, the symbols are experimental data points representing product concentration by HPLC, and the line is the dispersion model fit which gives vessel axial dispersion number (D/uL) and the actual mean residence time (τ). A description of how to calculate D/uL and τ from the experimental data is described in the Supporting Information and elsewhere. 10,2 The x-axis t/τ is the normalized time. The y-axis for startup transition curves is C/Cf, the normalized concentration, where Cf represents full strength steady state concentration after the transition. The y-axis for shutdown transition curves is C/Co, where Co represents full strength steady state concentration before the transition. D/uL was 0.0015 during startup transition (Figure 3) and D/uL was 0.001 during shutdown transition (Figure 4). This is very low axial dispersion. Also, from the startup transition, we verified that the reactor operated 94% liquid filled at baseline conditions. This is because flow was in the uphill direction through the coiled tubes and gas bubbled through the reactor faster than liquid flowed.


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1 0.8 0.6 0.4 0.2 0 0.7

0.8

0.9

1 t/τ

1.1

1.2

1.3

Figure 3. Startup transition and dispersion model fit for DARA reaction product in 780 mL coiled tube reactor. D/uL = 0.0015.

C/Co for product concentration by HPLC

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C/Cf for product concentration by HPLC

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1 0.8 0.6 0.4 0.2 0 0.7

0.8

0.9

1 t/τ

1.1

1.2

1.3

Figure 4. Shutdown transition and dispersion model fit for DARA reaction product in 780 mL coiled tube reactor. D/uL = 0.001.

Reaction Performance in Pipes-in-Series Reactors DARA reaction in 780 mL horizontal pipes in series continuous reactor. Pictures of the 780 mL horizontal pipes in series continuous reactor and oil heating bath are shown in the Supporting Information. The 316L stainless steel reactor was fabricated from 29 pipes that were each 9.53 mm o.d., 7.75 mm i.d., and 0.533 m long, connected by 29 smaller diameter jumper tubes that were 3.175 mm o.d., 1.75 mm i.d., and 0.7 m long each. Number of pipes in series



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was chosen to achieve about the same axial dispersion as the coiled tube reactor (D/uL ~ 0.001). A later section of this paper shows that a larger number of pipes in series results in lower overall axial dispersion number. The reactor achieved D/uL = 0.0007 during startup transition of the DARA reaction. The data and model fit are shown in the Supporting Information. The pipes were mounted on a slight uphill angle rather than exactly horizontal. The outlet end of each pipe was 1 inch higher than the inlet end. The jumpers were angled downhill to return to the bottom of the next pipe in series. This way, the reactor operated mostly liquid filled because the gas bubbled up the pipes faster than liquid flowed. From the startup transition, we verified that the reactor operated 94% liquid filled at baseline conditions. If we were building this reactor again, we would make the jumper tubes 1.59 mm o.d., 0.56 mm i.d., and 1 m long, to allow the reactor to run higher % liquid filled, and also to more completely re-saturate with H2 gas between pipes in series, although it would result in higher pressure drop. During the continuous flow experiment, step changes were made in τ (8, 10, 12 hours), temperature (100 °C, 115 °C, and 130 °C), and pressure (500, 1000, 1250 psig). Results are shown in Table 4. Table 4. Experimental results for continuous reaction in 780 ml horizontal pipes in series reactor. a Entry

τ

T (°C)

P (psig)

3 (%)

4 (%)

ee (%)

115

1000

91.7

5.9

-

1

(h) 10

2

12

115

1000

92.8

5.92

95.4

3

12

115

500

85

7.2

93.3

4

12

115

1250

90

6.8

96.9

5

12

115

1000

89.5

6.8

-

6

8

130

1000

87

7.3

94.1

7 8 100 1000 52 5.9 96.3 Reaction conditions were ketal 1 (1 eq), aminotetrazole 2 (1.1 eq), THF (7 volumes), CSA (0.02 eq), TBAI (0.01 eq), S:C 4000:1. a


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Entries 1 and 2 were from the same 40 hour continuous reaction experiment. Increasing τ from 10 hour to 12 hours increased conversion to product (3) from 91.7% to 92.8%, all else constant. This is almost identical to the batch reaction data in the well-mixed autoclave stirring at 315 rpm, which achieved 91.3% product (3) after 8 hours and 92.5% product (3) after 24 hours (Figure 2). As seen in Figure 2, first order rate curves were fitted to the experimental data. At the 12 hour time point for the higher stirring rate experiments, these curves pass through 92.4% product at 1000 psig (which is approximately the same as 92.8% product at 12 hours in the continuous pipes in series reactor). Conversion to ketone in the batch autoclave was about 7.5 to 8.3% after 24 hours, which was slightly higher than for the continuous reactions at 12 hours τ, which ranged from 5.9% to 7.3%. On-line HPLC data trends for this 40 hour continuous reaction experiment is shown in Figure 5. 94.00

area% desired product (3)

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Change to 12 hr τ

93.00

12 hr τ 92.00

10 hr τ Pushout of Samples held in reactor overnight

91.00

90.00 8

13

18

23

28

33

38

43

48

Time from start of continuous run (h)

Figure 5. On-line HPLC data trend for 40 hour continuous reaction experiment 1 in 780 ml horizontal pipes in series reactor.

Entries 3, 4, and 5 in Table 4 were all from the same 65 hour continuous reaction experiment. After 25 hours, pressure was changed from 500 psig to 1250 psig, which increased



conversion to product from 85% to 90%, all else constant. Then, after 50 hours, pressure was changed to 1000 psig, which dropped conversion to 89.5% product. Reaction performance at 500 psig was similar to batch. Referring again to Figure 2, the first order rate fit to the data at 500 psig and 315 rpm showed 86.4 % product (3) at 12 hours, which is similar to entry 3 in Table 4, 85% product (3) in 12 hours at the same pressure. All other process conditions were the same except reactor type. On-line HPLC data trend for this 65 hour continuous reaction experiment is shown in Figure 6. There was more scatter to the on-line HPLC data at the 500 psig conditions because of problems with the column, which was swapped after the first 28 hours.

The columns were not robust to a large number of injections on the LC method. The

first part of this run was done with an older column that had much more dispersion due to the aging process. As a result, the peaks were broader and lower intensity with smeared tails, and thus harder to integrate. 92.00 Area Percent Desired Product 3

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90.00 88.00 86.00

1225 psi 84.00

1000 psi 82.00

500 psi 80.00 0

10

20

30

40

50

60

70

Time from start of continuous run (h) Figure 6. On-line HPLC data trend for 65 hour continuous reaction experiment 2 in 780 ml horizontal pipes in series reactor.


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Entries 6 and 7 in Table 4 were from the same 40 hour continuous reaction experiment in the 780 mL horizontal pipes in series reactor. Making a temperature step change during the continuous run from 130 °C to 100 °C dropped % product (3) from 87% to 52%, all else constant. On-line HPLC data trends for this 40 hour continuous reaction experiment is shown in Figure 7. Between the 130 °C conditions and the 100 °C conditions, the feed pump was switched to pure THF for 7 hours, which is why you see the on-line HPLC data points at 0% for several hours in the transition. This was done to conserve reaction starting materials during the temperature adjustment and to get another check on axial dispersion in the middle of the flow experiment. 100.00

130 °C

90.00 Area Percent Desired Product 3

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


80.00 70.00

100 °C

60.00 50.00 40.00 30.00 20.00 10.00 0.00 0

5

10

15

20

25

30

35

40

Time from start of continuous run (h) Figure 7. On-line HPLC data trend for 40 hour continuous reaction experiment 3 in 780 ml horizontal pipes in series reactor.

Reagent and catalyst feed were held in the pumps for 11 days then run with no apparent deleterious results on conversion or impurity profile. This was done to verify feed solution and catalyst solution stability.



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Side by side comparison of coiled tube versus pipes in series. Two additional experiments were done with the 780 ml coiled tube reactor and 780 ml pipes in series reactor to get a comparison of the two reactor geometries at the same process conditions. Reaction operating parameters were S:C 4000, 8 hour τ, 1000 psi H2, and temperature only 100°C so that conversion would deliberately be only about 60%. Full conversion was not desired because we wanted to exaggerate differences in area% product at 8 h between the two reactors. Startup transitions confirmed that both reactors ran about 94% liquid filled at steady state. Substrate pump flow rate was 1.197 ml/min (0.416 mol/L) and catalyst pump was 0.0277 ml/min (0.0045 mol/L). Area % product (3) was about 52% in both cases. This confirmed that there was no apparent difference in reactor performance between the two continuous reactor geometries. This was confirmation that the pipes in series concept was valid. All further scale up work was done with pipes in series rather than coiled tubes. DARA reaction in 32 L horizontal pipes in series continuous reactor. The DARA reaction was next scaled up in a 32 L horizontal pipes in series continuous reactor. Pictures of the reactor and constant temperature heating bath are shown in the Supporting Information. The stainless steel reactor was fabricated from 29 pipes that were each 44.4 mm o.d., 26.9 mm i.d., and 1.86 m long, connected by 29 smaller diameter jumper tubes that were 6.35 mm o.d., 4.57 mm i.d., and 2.07 m long. The main reasons that the down flow jumper diameters were small were to minimize vapor space in the reactor and increase gas liquid mass transfer. The pipes were mounted on a slight uphill angle rather than exactly horizontal so that the reactor operated almost completely liquid filled, as the gas bubbled up the pipes faster than liquid flowed. The outlet end of each pipe was about 2 inches higher than the inlet end, and the outlet reducers were welded near the top of the circular end. If we were building this reactor again, we would make


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the jumper tubes 3.18 mm o.d., 1.75 mm i.d. to allow the reactor to run higher % liquid filled, and also to more completely re-saturate with H2 gas between pipes in series, although it would result in higher pressure drop. The entire reactor was submerged in a 500 L constant temperature oil bath maintained at 120 °C. The heating bath was filled with 105 gallons of Paratherm MG which was heated by submerged steam coils. The 73 hour continuous experiment began with seasoning run of reactor for 16 hours (10x dilute), followed by a switch to full strength reagent feed for 53 hours to react 15 kg ketal (1) starting material, followed by a solvent pushout for 14 hours. The procedure for the seasoning run is given in the experimental section. This was done 10X dilute compared to full strength. The purpose was to make sure that the metal surfaces were not hindering the catalyst and ligand, and to establish that the reaction was working properly at scale while committing minimum amount of substrate. During the campaign, step changes were made to τ (10, 12, 14 hours) and S:C (4000:1, 3500:1). Experimental results are shown in Table 5 and Figure 8. The best τ and S:C conditions were 14 hours and 3500:1 S:C loading, which resulted in 87.1% product (3), 94.3% ee, and 10.5% undesired ketone (4). Average pressure drop from reactor inlet to reactor outlet was about 5 to 10 psig. Pressure trends are shown in the Supporting Information. Table 5. Experimental results for DARA reaction 32 L horizontal pipes in series continuous reactor. a

Entry

τ (h)

S:C

3 (%)

4 (%)

%ee

1

12

4000

84.5

10.9

96.3

2

14

4000

85.7

10.9

96.3

3 14 3500 88 10.5 94.3 Reaction conditions were ketal 1 (1 eq), aminotetrazole 2 (1.1 eq), THF (7 volumes), CSA (0.02 eq), TBAI (0.01 eq), 3-4 equivalents H2, 120 °C, 1000 psig. a



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

On-line HPLC data for the 73 hour continuous run is shown in Figure 8. The figure only shows data during the full strength part of the run, from hour 20 to hour 73. Substrate pump flow rate was set at 41.48 ml/min (0.410 M solution) and catalyst pump flow rate was set at 0.141 ml/min (0.02965 M solution). Hydrogen was 3 to 4 molar equivalents through the campaign relative to ketal, and hydrogen feed rate was quantified from change in pressure on hydrogen feed cylinder as described in the Supporting Information. Area % product started out at only 84.5%. To improve conversion to product, τ was extended from 12 hours to 14 hours by changing substrate solution pumping rate to 35.6 ml/min, catalyst solution pumping rate to 0.12 ml/min, and H2 feed rate was decreased about 20%. Area % product only increased to 85.7%. Therefore, S:C was adjusted to 3500 by changing ketal solution pumping rate to 35.5 ml/min and catalyst solution pumping rate to 0.138 ml/min, resulting 88% desired product (3). Conversion to product (3) was lower than desired because of about 11% hydrolysis to ketone (4). The feeds were contaminated with water for this scale up experiment. Finally, to determine if 14 hours τ was necessary at this higher catalyst loading, τ was decreased to 10 hours to test impact of conversion. This was accomplished by changing substrate solution pumping rate to 49.7 ml/min and catalyst solution pumping rate to 0.193 ml/min. However, we could not reach steady state before running out of reagent feed, as seen in Figure 8. This campaign also demonstrated the value of on-line analytical, because the changes described above were all done based on the near real time analytical. The on-line HPLC can also confirm when steady state is not achieved.


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89.00

14 h τ, 3500:1

Transition to 10 h τ

88.00

Area % desired product (3)

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


87.00

14 h τ, 4000:1

86.00

85.00

Transition to THF only for solvent pushout

84.00

Transition to 14 h τ 12 h τ, 4000:1 83.00 20

25

30

35

40

45

50

55

60

65

70

Time from start of continuous run (h)

Figure 8. On-line HPLC data showing impact of and catalyst loading on conversion in 32 L horizontal pipes in series reactor.

The overall mass balance for the run was 99.0%. The reaction product solution was worked up and isolated as a batch. This resulted in 85% yield of isolated dried product that had 97% HPLC purity and 98.5% ee. These were excellent results despite the fact that conversion to undesired ketone (4) was high and area% product in the reactor was low. We attribute the lower area% product to water in the feed rather than reactor performance. We quantified several important parameters by measuring and modeling the transition curves during startup, concentration change, and shutdown. F-curves were measured by collecting samples from the reactor outlet and analyzing by HPLC after step changes. F-curves are described in the Supporting Information and elsewhere.11, 10 This was done for step change from solvent to seasoning concentration (10X dilute), from seasoning concentration to full concentration, and then from full concentration to solvent for pushout at the end. Actual



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

measured τ was 12.2 hours at full strength reaction conditions. Axial dispersion number for this reaction process was 0.0008 when the step change was made from solvent to seasoning concentration, 0.0035 during the step change from seasoning concentration to full concentration, and 0.006 when step change was made from full concentration to solvent for pushout at the end. These are all low dispersion numbers. RTD was more broad when switching from concentrated to dilute conditions. It is important to quantify how much higher is the axial dispersion number for pushout F-curve compared to push-in F-curve when quantifying lot genealogy and deviation boundaries. The measured and model-fitted F-curves are shown in the Supporting Information. Results of the model fits are listed in Table 6. Table 6. Actual τ in the continuous reactor, and axial dispersion number, calculated by modeling transition curves. Entry

Initial Material

Final Material

τ actual (hr)

D/uL

1

THF

Seasoning

11.6

0.0008

2

Seasoning

Reagent

12.2

0.0035

3

Reagent

THF

9.5

0.0060

DARA reaction 11 L vertical pipes in series continuous reactor. Compared to the horizontal pipes reactor, it was hypothesized that the vertical pipes in series reactor design would have higher gas/liquid mass transfer rates, achieve low axial dispersion with fewer pipes in series because of better mixing in the radial direction, operate with higher % liquid filled, and occupy less plant footprint. Pictures of the continuous reactor and the steam heat tracing on each pipe are shown in the Supporting Information. Steam heat tracing was chosen because the hot oil bath was not practical in the vertical orientation. The pilot scale reactor carts are on wheels and portable, and a tall hot oil bath would be too top heavy and potentially tip over when rolled in and out of the laboratory bunker. Steam heat tracing was a safer option for the portable equipment. The 316L stainless steel reactor was fabricated from 15 pipes that were each 25.4 mm o.d., 22.1 mm i.d., and 1.83 m tall, connected by 15 smaller diameter jumper tubes that were


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6.35 mm o.d., 4.57 mm i.d., and 2.06 m long. If we were building this reactor again, we would make the jumper tubes 3.18 mm o.d., 1.75 mm i.d. to minimize vapor space in the reactor, increase gas liquid mass transfer, and minimize the potential for surging, although it would result in higher pressure drop. Total amount of ketal starting material processed was 5 kg in the 98 hour continuous flow experiment. The startup was a seasoning run for 18 hours. Then, full concentration reagent feed solutions were pumped into the reactor for 62 hours. Finally, the shutdown transition with solvent pushout was 18 hours. τ and H2 equivalents were deliberately changed during the campaign to test their impact on conversion during the 62 hour full strength portion (Table 7). Baseline conditions were 14 hours τ, 3.6 molar equivalents H2, 120 °C, 1000 psig, 3500:1 S:C, and feed compositions as indicated in the footnote to Table 7. Decreasing τ from 14 hours to 10 hours decreased area % product from 87.5% to 85.7%. Decreasing H2 molar flow rate from 3.6 eq to 2.3 eq did not have significant effect on conversion. These transitions can be seen in Figure 9, which shows on-line HPLC data for conversion as a function of τ and equivalents H2 during the continuous production run in the 11 L reactor. Average pressure drop from reactor inlet to reactor outlet was about 40 psig. It was not surprizing that overall reactor pressure drop was higher for vertical pipes in series reactor versus horizontal pipes reactor. Pressure trends are shown in the Supporting Information. Table 7. Experimental results for DARA reaction 11 L vertical pipes in series continuous reactor. a

Entry

τ (h)

Eq H2

3 (%)

4 (%)

1

13.3

3.6

87.5

10.8

2

10

3.6

85.7

10.7

3 14 2.3 87.5 10.6 Reaction conditions were ketal 1 (1 eq), aminotetrazole 2 (1.1 eq), THF (7 volumes), CSA (0.02 eq), TBAI (0.01 eq), S/C 3500, 120 °C, 1000 psig. a



88.5

Condition 1: 13.3 h τ, 3.6 eq H2

Condition 3: 14 h τ, 2.3 eq H2

88 area% desired product (3)

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

87.5 87 86.5 86 85.5 85 11.00

Transition from seasoning 21.00

31.00

Condition 2: 10 h τ, 3.6 eq H2 41.00

51.00

61.00

Time from start of continuous run (h) Figure 9. On-line HPLC data showing impact of τ and eq H2 on conversion in 11 L vertical pipes in series reactor.

Part of the reaction product solution was forward processed by batch workup and isolation. 1.6 kg of the product was isolated in about 80% isolated yield, with 97.9% purity and 98.68 %ee. Once again, like with the 32 L horizontal pipes reactor experiment, conversion to undesired ketone 4 was high which made conversion to desired product 3 lower. We believe this was more due to known high water in the feed rather than the reactor performance. Actual measured τ was 13.3 hours at full strength reaction conditions, determined by measuring and modeling the transition curves during startup concentration change. Axial dispersion number for this reactor was 0.004 when the step change was made from solvent to seasoning concentration, 0.004 during the step change from seasoning concentration to full concentration, and 0.012 when step change was made from full concentration to solvent for pushout at the end. These are all low dispersion numbers. Not surprisingly, RTD is more broad when switching from concentrated to dilute conditions. The measured and model-fitted F-curves are shown in the Supporting Information. Results of the model fits are listed in Table 8. This


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dispersion number was higher than the horizontal pipes reactor because the reactor only had 15 pipes in series rather than 29. The impact of number of pipes in series on dispersion number is described later in this paper. Nevertheless, only 15 vertical pipes in series were needed to achieve the goal of D/uL less than or equal to 0.01. Table 8. Actual τ in the 11 L continuous reactor, and axial dispersion number, calculated by modeling transition curves. Entry

Initial Material

Final Material

τ actual (hr)

D/uL

1

THF

Seasoning

13.4

0.004

2

Seasoning

Reagent

13.3

0.004

3

Reagent

THF

14.7

0.012

Ru catalyzed DARA reaction in 410 mL vertical pipes in series continuous reactor. A picture of a 410 ml vertical pipes in series reactor and the vessel used for hot oil bath is shown in the Supporting Information. The reactor was fabricated from 15 stainless steel pipes (316L) that were each 9.53 mm o.d., 7.75 mm i.d., and 0.579 m long, connected by 15 smaller diameter jumper tubes that were 1.59 mm o.d., 0.56 mm i.d., and 0.64 m long each. The continuous DARA reaction in the 410 ml reactor used Ru catalyst instead of Ir, as shown in the following scheme. Comparison of flow reactor types would be better if the Ir catalyst system had been run in the 410 ml vertical pipes in series reactor instead, but this reactor was only used with the Ru chemistry, nevertheless it is included because it was compared directly to the batch. The main reason for switching to ruthenium was because we could telescope steps (the Ir process required additional isolation) but this is beyond the scope of this paper. Notably, the reaction solvent was 97% MeOH and 3% THF, instead of 100% THF. The 3% THF was required so that reagent mixture remains homogeneous at room temperature (20 °C). The reaction was conducted at 120 °C, 1000 psig, 1.5 molar equivalents H2, with S:C 1000:1, and ketal (1) concentration 0.42 M in combined reagent and catalyst solutions. The



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product (3) was soluble at reaction temperature, but would precipitate at room temperature and could potentially foul at the outlet of the reactor. To prevent this, 1/1 volume ratio of toluene was mixed with the product solution continuously in a mixing Tee at the reactor exit , submerged in the same hot oil bath. At the 8 hour τ, reagent solution pump flow rate was 0.749 ml/min, catalyst was 0.04139 ml/min, and toluene at reactor exit in the heated zone was 0.791 ml/min.

Scheme 2. Ruthenium Catalyzed DARA Reaction

The continuous reactor operated almost completely liquid filled because the gas bubbled up the pipes faster than liquid flowed. Accounting for thermal expansion of the methanol, the reactor ran about 98% liquid filled, which was higher than the horizontal pipes in series reactor or coiled tube reactor which both operated 94% liquid filled. This was because of the smaller diameter jumper tubes, the lower H2 equivalents, and the fact that bubbles move through the pipes at higher velocity in the vertical orientation. The continuous reactor was run with 6, 8, and 10 hours τ. Three samples were taken at each τ, each 1 hour apart, to prove steady state. Results are shown in Table 9. Area % product (3) was about the same at all 3 τ (95.6% to 96.1%). Ketone (4) was lower at the 10 hour τ, about 0.7% versus 1.9%. We cannot make conclusions about rate of conversion to product (3) in the continuous reactor because we did not take samples


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at shorter τ. Compared to the Ir catalyzed reaction in THF, conversion to ketone (4) was lower (about 2% versus 7-10%) but ee was also lower (about 93% versus about 95%). Table 9. Experimental results for continuous reaction in 410 ml vertical pipes in series reactor.

τ (hours) 6

8

10

1 (%)

3 (%)

4 (%)

ee (%)

0.15

95.59

1.89

93.4

0.15

95.57

1.91

93.4

0.14

95.58

1.89

93.4

0.15

95.09

2.01

91.9

0.14

95.41

1.68

92.1

0.14

95.51

1.58

92.0

0.09

96.14

0.76

93.0

0.08

96.13

0.7

93.0

0.08

96.16

0.67

93.1

The reaction was run batch and samples were taken over time from 2 to 24 hours to characterize reaction rate. Results are shown in Table 10. At this catalyst loading, reaction reached 93.7% product (3) after just 2 hours, and 96.2% after 4 hours. Final reaction results were about the same batch and continuous, but we would need more data points in the continuous reactor at shorter τ to compare reaction rate. Table 10. Batch reaction results for DARA using Ru catalyst. a

a

Time

1 (%)

3 (%)

4 (%)

0

41

21.39

24.9

2

1.6

93.7

3.96

4

0.49

96.2

2.43

5.5

0.24

97

1.78

8

0.09

97.5

1.23

10

0.08

97.7

1.01

24 0.09 98.1 0.41 Temperature was 120 °C, pressure was 1000 psig, S:C was 1000:1, and CSA was 2 mol %.

As with the other types of continuous reactors, axial dispersion was quantified experimentally by sampling during startup transition and shutdown transition by analyzing for concentration of product in solution. Experimental data and model fits are shown in the



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Supporting Information. Dispersion number D/uL was 0.002 during startup transition and 0.009 during shutdown transition. This is low axial dispersion, but not as low as the 29 horizontal pipes in series, which is not surprising. Another section in this paper shows that decreasing number of pipes in series increases D/uL. Nevertheless, only 15 vertical pipes in series were needed to achieve the goal of D/uL less than or equal to 0.01. Characterizing reactors using gas/water flow in transparent tubes and pipes. Vapor/liquid mass transfer. The experimental results in this paper proved that gas/liquid mixing rate was sufficient in all 3 types of continuous reactors (coiled tube, horizontal pipes in series, and vertical pipes in series) for the DARA reaction. Reaction performance in the flow reactors was about the same as in well-mixed batch. Simple mixing Tees were sufficient for initial gas + liquid contacting at the reactor inlet. No special diffusers were necessary at the inlet to each pipe in series. While gas/liquid mixing in each continuous reactor was sufficient, we still wanted to differentiate between reactor type and scale, and to quantify kLa in the tubing jumpers, which would require more direct measurement. kLa is described in the equation dC/dt = kLa (C*C), where C is the dissolved gas concentration and C* is the equilibrium dissolved gas concentration. It is a lumped parameter that includes resistances to mass transfer on both the gas and liquid sides of the interface as well as the total interfacial area per unit volume, and thus has units of 1/time. The higher the kLa relative to reaction rate, the higher the dissolved H2 concentration available for reaction, with the maximum being equilibrium saturation. It has been quantified by others in small scale continuous reactors, such as bubble columns,12 falling film micro reactors,13 microchannel devices ,14 In those studies kLa values ranged from about 0.1 s-1 to 10 s-1. It is difficult to directly measure dissolved hydrogen concentration in the solvent. As an alternative, we measured the rate of oxygen dissolving into water from air at 60 psig and 22


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°C. Pictures of 2-phase water-air flow through PFA tubing of different diameters, vertical 1” pipe, and horizontal 1” pipe, are shown in Figure 10 through Figure 15.

Figure 10. Gas and liquid segmented flow in 0.25 inch o.d., 0.125 inch i.d. tubing; air and water at 60 psig, 20 °C, and 40 ml/min liquid flow.





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

Figure 13. Liquid and gas slug flow in 0.75 inch o.d., 0.625 inch i.d. “horizontal” coiled tube (2% incline); air and water at 60 psig, 20 °C, and 40 ml/min liquid flow. Flow direction from left to right.

Figure 14. Liquid and gas bubble flow in 1 inch i.d. vertical pipe; air and water at 60 psig, 20 °C, and 40 ml/min liquid flow.

Figure 15. Liquid and gas slug flow in 1 inch i.d. “horizontal” pipe (2% incline); air and water at 60 psig, 20 °C, and 40 ml/min liquid flow. Flow direction from right to left.

An on-line Ocean Optics® HIOXY T1000-TS-NEO dissolved oxygen probe measured the oxygen concentration in water. The water was initially sparged with N2 for several hours so that dissolved oxygen concentration was approximately zero (below detection limit). The results for kLa in 0.25, 0.125, and 0.0625 inch o.d. PFA tubing are shown in Figure 16 and Figure 17. The water and air mixed in a simple mixing T, flowed through a length of the tubing, and into an


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overflow Tee where the dissolved oxygen probe measurement was made. Multiple lengths of each size tubing were measured, and multiple repeat data points were collected for each condition. Therefore a number of data points were collected to calculate each kLa data point. In the pipes in series reactors, this represents the tubing that is used for the inlet to the first pipe, and the jumpers between pipes in series. The pipes in series reactors are designed for high kLa in the inlet and jumper tubing so that the liquid will re-saturate with gas between each pipe. Several observations can be made about kLa from the data in Figure 16 and Figure 17. 1. kLa greater than 0.2 s-1 is achievable for gas/liquid segmented flow in the jumper tubing at all scales. Research scale flow rates are about 1 to 5 ml/min, pilot scale flow rates are about 10 to 40 ml/min, and manufacturing scale flow rates are 200 ml/min or greater. For 1/1 gas/liquid volumetric flow ratio, kLa larger than 0.2 s-1 in the inlet and jumper tubes is achievable at all scales, depending on the choice of tubing diameter (Figure 17). 2. kLa increases with linear velocity in the jumper tubing. In all 6 cases shown in the data in Figure 16, kLa increases with increasing flow rates in the tubing, approximately linearly. The disadvantage of running at higher liner velocities is the overall pressure drop across the reactor increases. In general, we would recommend 0.0625 inch o.d. tubing for research scale continuous reactors with flow rate below 5 ml/min, 0.125 inch o.d. tubing for pilot scale reactors with flow rate in the range 5 to 100 ml/min, and 0.25 inch o.d. tubing for manufacturing scale reactors with flow rate in the range 100 to 200 ml/min (Figure 17). If flow rate is higher than 200 ml/min, larger diameter tubing may be considered. There will be a trade-off between vapor liquid mass transfer rate and pressure drop in the tubing. 3. kLa increases at higher gas/liquid volumetric flow ratios. In all 3 sizes of tubing and for all linear velocities tested, when gas/liquid volumetric flow ratio increased from 0.2 to 1, kLa



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approximately doubled because of higher interfacial surface area and higher linear velocities. Higher gas flow rates increase gas/liquid mixing, but also increase gas waste (excess equivalents) and increase pressure drop through the reactor. 4. kLa can be maintained constant in the reactor inlet tubing and jumpers with scale-up/scaledown. As can be seen in the trends in Figure 16, kLa is a strong function of the choice of tubing size, for a given liquid and gas flow rate. Suppose you wanted to achieve kLa greater than 0.1 s-1 in the tubing jumpers at research scale, pilot scale and manufacturing scale. This is possible, but it impacts the suitable size of jumper tubing diameter. 5. kLa in the pipes was much lower than in the jumpers. kLa in the vertical pipes was on the order of 0.001 to 0.01 s-1. For example, consider data in Figure 18 for the case of 200 ml/min water and 1/1 gas/liquid flow ratio. Inlet to the pipe was 68% of saturation, outlet was 96% of saturation, and mean residence time was 5 minutes, which corresponds to a kLa about 0.007. The kLa quantities stated above are for O2 dissolving into water at room temperature and 60 psig, not H2 dissolving into THF or methanol at 120 °C and 1000 psig. More testing is needed for H2 dissolving into the solvents at the elevated temperatures and pressures. Much higher kLa values can be found in the literature for gas/liquid reactions in microstructured PFRs. For example kLa was on the order of 0.8 s-1 to 2.2 s-1 for ozonolysis reactions in Corning glass micro reactors with liquid flow rates in the range 1 to 10 ml/min.15 However, such high kLa values are simply not necessary for asymmetric hydrogenation reactions with 8 to 14 hour reaction time which are limited by catalyst turnover rate at the low catalyst loadings.


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1.4

1.2

1/4", 1/1 gas/liquid 1/4", 0.2/1 gas/liquid 1/8", 1/1 gas/liquid

1

1/8", 0.2/1 gas/liquid 1/16", 1/1 gas/liquid

kLa (1/s)

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


0.8

1/16", 0.2/1 gas/liquid

0.6

0.4

0.2

0 0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

Linear Velocity (m/s) Figure 16. kLa as a function of linear velocity and tube diameter measured for oxygen uptake into water for 2-phase segmented gas-liquid flow in standard tubing.



1.4 1/4", 1/1 gas/liquid 1/4", 0.2/1 gas/liquid 1/8", 1/1 gas/liquid 1/8", 0.2/1 gas/liquid 1/16", 1/1 gas/liquid 1/16", 0.2/1 gas/liquid

1.2

1

kLa (1/s)

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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0.8

0.6

0.4

0.2

0 0

20

40

60

80

100

120

140

160

180

200

Liquid Flow (mL/min) Figure 17. kLa as a function of liquid volumetric flow rate and tube diameter measured for oxygen uptake into water for 2-phase segmented gas-liquid flow in standard tubing.

It was difficult to measure how much gas uptake occurred in the pipes, because so much gas uptake happened in the small diameter tube between the mixing Tee and the inlet to the pipe, and then in the tube from the outlet of the pipe to the Ocean Optics® probe. However, a comparison was made between a 1 L horizontal pipe, and 1 L vertical pipe, and a 1 L section of 0.75 inch o.d. 0.625 inch i.d. coiled tubing. The 0.625” i.d. coiled tubing represented the type of continuous hydrogenation reactor that we have used historically. 1 Data is shown in Figure 18. Vapor/liquid mass transfer was best for the vertical pipes in series. Figure 18 shows 4 sets of bar graphs, each corresponding to a different liquid flow rate and gas/liquid flow ratio. In all 4 cases, the bar on the left of each set represented % saturation at the inlet to the 1 inch pipe or


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0.75 inch tube, and the next three bars represent % saturation at the outlet. Relatively speaking, the vertical pipe orientation was best for gas/liquid mass transfer. The 0.75 inch coiled tube was better than the horizontal pipe for achieving high % saturation, but the 0.75 inch coiled tube reactors were not constructed with frequent small diameter tubing jumpers, which is where vapor/liquid mass transfer rate is highest. Therefore the pipes in series reactor design with small diameter tubing jumpers is a better design for gas/liquid mass transfer, even if the pipes are oriented horizontally. Accordingly, constructing the reactors with a larger number of pipes in series can maintain higher concentration of dissolved reagent gas throughout. An attractive option is to reduce the diameter (and thus τ) of the first several pipes in series where gas uptake is higher because of higher reaction rate. 10 ft ¼” tube 10 ft ¼” tube + horizontal pipe 10 ft ¼” tube + ¾” tube coil 10 ft ¼” tube + vertical pipe 100% 90% 80% Percent of full saturation

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


70% 60% 50% 40% 30% 20% 10% 0% 200 mL/min, 0.2/1 gas/liquid

200 mL/min, 1/1 gas/liquid

40 mL/min, 0.2/1 gas/liquid

40 mL/min, 1/1 gas/liquid

Figure 18. Gas uptake in coiled tubes versus horizontal pipes versus vertical pipes. Pipe was 1 inch i.d. and 6.1 feet long, made from clear PVC. 0.75 inch tubing was 0.625 inch i.d. and was coiled in slightly uphill orientation so that exit end was 1 inch higher than inlet end.



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RTD. Another important question was which continuous reactor type is best for minimizing axial dispersion (i.e. which is closest to ideal plug flow). The goal was D/uL < 0.01. All types of reactors reported in this manuscript met this criteria, both at research and pilot scale. F-curves for startup and shutdown transitions are shown for the coiled tube reactor in Figure 3 and Figure 4 and the rest are shown in the Supporting Information. Table 11 summarizes all of the dispersion testing results. For all reactor types at both research scale and pilot scale, D/uL ranged from 0.0008 to 0.012. Practically speaking, this means that steady-state concentrations and conversion was achieved after only 1.1 to 1.3 volume turnovers during start up transition, and that reactor contents were 99.9% pushed out after 1.1 to 1.4 volume turnovers during shutdown transitions. There was no significant impact of scale up on dispersion because all reactors were designed for low dispersion. Table 11. Axial dispersion number as a function of reactor type and scale. D/uL for reaction startup transition 780 ml coiled tube reactor 0.0015 780 ml horizontal pipes reactor, 29 pipes 0.0007 410ml vertical pipes reactor, 15 pipes 0.002 32 L horizontal pipes reactor, 29 pipes 0.0008a, 0.0035b 11 L vertical pipes reactor, 15 pipes 0.004 a, 0.004 b a Startup transition from solvent to seasoning reaction concentration (10X dilute). b Startup transition from seasoning concentration to full strength.

D/uL for reaction shutdown transition 0.001 0.009 0.006 0.012

Data in Table 11 does not offer a fair comparison of horizontal pipes versus vertical pipes in series because there is a larger number of horizontal pipes in series. Therefore an experiment was designed to directly investigate D/uL in the horizontal versus vertical orientation. The test was done with water flowing through two 1 inch i.d., 6 ft tall, pipes in series (Figure 19). The data in the figure shows typical C-curves, which are the response to a pulse tracer injection. This pulse tracer was sodium chloride and it was measured by on-line refractive index probe at the exit from the pipes. In this test, the vertical orientation was superior for minimizing axial dispersion. D/uL was 0.31 when the 2 pipes were mounted horizontally and 0.13 when the 2


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pipes were mounted vertically, all else constant. This is because bubble flow in the vertical tubes increased mixing in the radial direction more than it did in the horizontal orientation. Looking at Figure 14 and Figure 15 helps to explain why. The gas phase separated to the top of the horizontal pipe, whereas the gas bubbles were more evenly distributed in the radial direction for the vertical pipes.

Weight % NaCl

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


0.5 0.45 0.4 0.35 0.3 0.25 0.2 0.15 0.1 0.05 0

Horizontal D/ul = 0.3064 Vertical, D/uL = 0.1346

0

1000

2000

3000

Time (s) Figure 19. Impact of horizontal versus vertical orientation on axial dispersion.

Next we investigated the impact of number of pipes in series on D/uL (Table 12). The flow test was done with 1 inch i.d., 6 foot tall transparent pipes, water flow, and RI detection of dissolved NaCl. It is clear from the data that D/uL decreases when number of pipes increases. This is not surprising. There is no back-mixing from one pipe to the previous. Even if axial dispersion were infinity in each pipe, the RTD would not be worse than that of the same number of CSTRs in series. Table 12. Effect of number of vertical pipes in series on dispersion number. Number of Pipes 29 15 4 2

Dispersion number, D/uL 0.0015 0.0065 0.0525 0.1032



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Reactor pressure drop. Pressure drop was highest for the vertical pipes in series reactor design and lowest for the coiled tube design. Continuous high-pressure hydrogenation reactions were run in three different pilot scale reactors. Reactor delta P was 40 psig for an 11 L vertical pipes in series reactor, 7 psig for a 32 L horizontal pipes, and 3 psig for a 32 L coiled tube PFR. Dimensions of the pipe and tubing lengths, diameters, and the number of pipes or coils in series are listed in Table 13. Reactor ∆P was the measured difference between inlet pressure and outlet pressure and steady-state flow conditions. Experimental results are shown in Table 13. The vertical pipes in series reactor had higher pressure drop than the horizontal, even though it only had 15 pipes in series rather than 29, the volume was about 3X smaller, and the flow rates were lower. Overall pressure drop limits the maximum throughput, because ∆P increases when flow rates increase. The horizontal pipes in series reactors can be used over a much wider flow rate range because pressure drop is lower. The coiled tube reactor had the lowest pressure drop overall, despite the fact that temperature was lower and thus viscosity was higher. The 32 L coiled tube reactor campaign was for a different unpublished chemistry. A picture of the 32 L coiled tube reactor is shown in the Supporting Information. Table 13. Reactor pressure drop for different types of pilot scale continuous reactors. Reactor

Description

540 feet of tubing 0.62 inch i.d., 5 coils in series connected by 4 ft long 0.18 inch i.d. 32 L coiled tube jumpers 29 pipes in series, each 6.1 feet long and 1.06 32 L horizontal inch i.d., slight uphill angle, connected by 0.18 pipes in series inch i.d. jumpers 15 pipes in series, each 6 ft long and 0.87 inch 11 L vertical i.d., mounted vertically, connected by 0.18 pipes in series inch i.d. jumpers a chemistry from a separate study not described in this paper.

Continuous Chemistry Campaign asymmetric H2 in 7 volumes THF, 1000 psig, 30°C, 0.5/1 gas/liquid ratio. a DARA in 7 volumes THF, 1000 psig, 115 °C, 0.5/1 gas/liquid ratio. DARA in 7 volumes THF, 1000 psig, 115 °C, 0.5/1 gas/liquid ratio.

Reactor ∆P

3 psig

7 psig

40 psig

Surging. To maintain segmented flow and prevent surging in the vertical pipes in series reactor, the down jumper tubes should be designed with small enough internal diameter to give a


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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


Froude number greater than 0.31 for a given gas flow, liquid flow, gas density, liquid density. 16 Froude number in the down jumper tubes is calculated from Equation 1. =

∙

∙

−

=liquid velocity =gravitational constant =internal diameter ρ=density of liquid or vapor Equation 1: Froude number

Surging can be an issue in the vertical pipes in series because of the influence of gravity on the downflow tubes. Surging was observed in the 14 L and 200 L vertical pipes in series transparent model reactors. Data is shown in Table 14. These model reactors were made with PVC transparent pipes, useful for observing bubble size, flow regime, axial dispersion, and surging. Pictures of these 14 L and 200 L transparent reactors are shown in the Supporting Information. Originally the downflow tubing jumpers were 0.18 inch i.d. on the 14 L reactor and 0.4 inch i.d. the 200 L reactor. They were tested with water and nitrogen flow, 20 to 100 psig inlet pressure, 12 hour liquid τ, and 1/1 gas/liquid volumetric ratio. Fr was 0.2 in the 14L reactor and 0.4 in the 200 L reactor. Significant surging was observed about once every 10 minutes, even though Fr in the 200 L reactor was greater than 0.31. The downflow tubing jumpers were changed to 0.069 inch i.d. on the 14 L reactor and 0.18 inch i.d. the 200 L reactor, changing Fr to 2.6 in both, and this eliminated the surging (Table 14). Table 14. Froude numbers calculated for flow scenarios in vertical pipes in series reactors.

τ (h)

liquid flow (ml/min)

downjumper i.d. (inches)

Fr

Surging observed

200 L clear pipes

12

278

0.4

0.4

Yes

200 L clear pipes

12

278

0.18

2.6

No



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

14 L clear pipes

12

19.4

0.18

0.2

Yes

14 L clear pipes

12

19.4

0.0625

2.6

No

11 L reactor

10

18.3

0.18

0.2

ND*

11 L reactor

14

13.1

0.18

0.1

ND*

11 L reactor

10

18.3

0.069

1.9

ND*

11 L reactor 14 13.1 *cannot observe because not transparent.

0.069

1.4

ND*

The 11 L vertical pipes in series reactor used in this study had 0.18 inch i.d. downjumpers. This resulted in Fr = 0.1 at reaction conditions. This is insufficient to prevent surging. Regardless, there were no deleterious effects because the expansion chambers in series pressure/flow control system was designed to tolerate surging, as described in the Supporting Information. The horizontal pipes in series reactors can be used over a much wider flow rate range because surging is not an issue, even when flow rates are low. Pressure control and gas flow control. Two different methods have been used for control of reactor pressure, gas flow, de-pressurization at the reactor exit, and vapor/liquid separation. These are described in detail in the Supporting Information. The main difference is whether or not the system includes a back pressure regulator at the reactor exit. When no back pressure regulator is used, the pressure in the reactor oscillates, by design. When a back pressure regulator is used, the reactor pressure remains constant. When no back pressure regulator is used, H2 gas mass flow is controlled indirectly on the back end of the reactor by the sequence time and volume of a series of expansion chambers. In contrast, when a back pressure regulator is used, H2 gas mass flow is controlled directly at the inlet of the reactor with a mass flow controller. The main benefit of not using the back pressure regulator is that the reactor system can handle some solids in flow without plugging or clogging. The main benefit of using the back pressure regulator is that the reactor pressure does not oscillate.


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Scale up and scale down of the pipes in series. Table 15 lists several of the continuous pipes in series reactors that have been used at Eli Lilly and Company for high pressure hydrogenations. As seen in the table, the reactors have been scaled down to 11 ml and scaled up to 360 L. 3 Pictures of the 11 mL and the 360 L reactors are shown in the Supporting Information. The last column shows % of reactor volume in the pipes (uphill flow direction). In general, the desire is to keep more than 98% of the total volume in the upflow pipes so that the reactor will run almost completely liquid filled even at high gas volumetric flow rate. However, for the smallest scale research reactors, less of the total volume is in the larger diameter pipes because of the desire for longer jumpers to achieve better gas/liquid mass transfer. The 360 L reactor is in GMP manufacturing. The 360L reactor, hydrogen supply, vapor liquid separation, and H2 stripping systems were all outside the plant building which was a key safety benefit to this reactor type. 3 Table 15. Pipes in series reactors used for high pressure hydrogenations.

Entry

Reactor volume (ml)

1

10.9

2

c

24

3

b

168

4

a

778

5

a

412

6

b

1235

7

2053

8

a

11,000

9

a

31,700

upflow pipes Jumpers upflow pipes Jumpers upflow pipes Jumpers upflow pipes Jumpers upflow pipes Jumpers upflow pipes Jumpers upflow pipes Jumpers upflow pipes Jumpers upflow pipes Jumpers

tubing o.d. (mm) 9.53 1.59 9.53 1.59 9.53 1.59 9.53 3.18 9.53 1.59 9.53 1.59 12.7 3.18 25.4 6.35 44.5 6.35

Tubing i.d. (mm) 7.75 0.56 7.75 0.56 7.75 0.559 7.75 1.75 7.75 0.56 7.75 0.559 10.9 1.75 22.1 4.57 26.924 4.57

length each tube (m) 0.0064 1 0.0064 3.66 0.0762 0.61 0.533 0.701 0.579 0.64 0.579 0.64 1.07 1.13 1.83 2.06 1.86 2.07

volume each tube (ml) 0.30 0.25 0.30 0.9 3.59 0.15 25.1 1.69 27.3 0.16 27.3 0.16 99.9 2.72 701 33.8 1060 34


Number in series 20 20 20 20 45 45 29 29 15 15 45 45 20 20 15 15 29 29

Total volume all tubes (ml) 6.03 4.9 6.03 17.9 162 6.73 729 49.1 409 2.35 1230 7.06 1999 54.4 10,500 507 30,700 987

Volume % of reactor 55.2 44.8 25.2 74.8 96 4 93.7 6.3 99.4 0.572 99.4 0.572 97.4 2.65 95.4 4.59 96.9 3.11


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

upflow pipes 25.4 22.1 1.83 701 45 31,600 99.3 Jumpers 3.18 1.75 2.13 5.15 45 232 0.729 upflow pipes 60.3 52.5 3.69 7989 45 359,500 99.1 b 11 363,000 Jumpers 6.35 4.57 4.27 70 45 3,150 0.869 3 c b a Reactors used for DARA reaction in this study. Reactors used for a reductive amination. Reactor used for 10

b

31,800

asymmetric hydroformylation. 2

Conclusions A high pressure, high temperature DARA reaction with 8 to 14 hours reaction time was run in three different continuous reactor types; coiled tubes, horizontal pipes in series, and vertical pipes in series. The pipes in series reactors were designed and developed as an alternative to the coiled tubes because they will be more scalable to large reactor volumes, for example the 360 L GMP manufacturing reactor shown in the SI, and ultimately to larger volumes. Baseline conditions for the DARA reaction were 1000 psig, 115 °C, Ir catalyzed with 4000:1 S:C, 1.1 eq aminotetrazole relative to ketal, 0.02 eq CSA, 0.01 eq TBAI, 7 volumes THF, 2.3 to 4 molar equivalents H2, and 12 hours reaction time. The reaction achieved about 90% conversion to desired product and 95% ee, in batch and in all 3 types of continuous reactors. The conditions were first developed batch, and batch kinetics showed that it was about a 12 hour reaction at baseline conditions. First order rate constant was about 0.5 h-1 with 1000 psig H2 at 4000:1 S:C, therefore this reaction was not a candidate for a small continuous reactor with short

τ. On-line HPLC was a useful tool to guide impact of process step changes in τ, pressure, and temperature and to monitor transitions from one steady state to the next. Compared to the coiled tube reactors, both the horizontal and vertical pipes in series reactors were more scalable and had higher kLa, but they had higher overall pressure drop. The vertical pipes in series reactors had the highest kLa, operated with highest % liquid filled, lowest axial dispersion for a given number of pipes in series, and smallest plant footprint, but had higher pressure drop and surging potential. In comparison, the horizontal pipes reactors have the ability to operate over a wider


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throughput range than vertical, because pressure drop is less of an issue at high flow rates and surging is less of an issue at low flow rates. Maintaining Froude number higher than 0.3 in the vertical pipes helps to prevent surging. In the vertical pipes in series reactor, full conversion was demonstrated with only 2.3 molar equivalents H2 in flow at pilot scale and 1.5 equivalents H2 in flow at research scale. Not less than 3 molar equivalents H2 were used in the production run with the horizontal pipes in series, therefore the impact of less H2 flow is unknown. Vapor liquid mass transfer rates were measured in the tubing jumpers with water and air at ambient temperature. kLa was not measured for hydrogen dissolving into THF solvent, but the surrogate testing with air and water was sufficient to compare the continuous reactors relative to each other. kLa values greater than 0.1 s-1 were achieved at all scales in the small diameter tubing, which is used at the reactor inlet and in the jumpers between each of the pipes in series. kLa was only on the order of 0.001-0.01 in the pipes themselves, but the strategy was to re-saturate with gas in the jumper tubes between each vertical pipe. kLa value was high enough for the 8 to 14 hour reaction so that the reaction was not mass transfer limited, as confirmed by conversion rates matching those of well-mixed batch reactors. Axial dispersion was low in all 3 types of continuous reactors at all scales, with D/uL ranging from 0.0008 to 0.012. A greater number of pipes in series gave lower axial dispersion. A greater number of pipes in series also results in higher overall kLa because liquid vapor mixing is highest in the jumper tubes. Two different methods of pressure and gas flow control were presented, one with a back pressure regulator and one without. The main benefits of not using the back pressure regulator is that the reactor system can handle some solids in flow without plugging or clogging, and a gas mass flow controller is not needed. The main benefit of using the back pressure regulator is that the reactor pressure control is more constant.



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

Experimental section Preparation of catalyst solution for 32 L reactor seasoning run. Catalyst solution was prepared in a glove box to avoid oxygen. Table 16 shows the materials used. Table 16. Catalyst solution preparation for seasoning run in 32 L continuous reactor. Material

Source

MW

Amount

[Ir(cod)Cl]2

UMICORE

671.71

0.327 g

(S)-Xyl-BINAP

Takasago

734.9

0.716 g

THF

Superior Solvents and Chemical, Indianapolis

72.11

0.140 L

A 300 ml glass pressure bottle and associated pressure head was rinsed with methanol. A clean stir bar was inserted, and the sealed bottle was purged with nitrogen until dry. The clean and dried pressure bottle with stir bar was weighed and then transferred into the glove box with an extra cap. About 200 mL of THF in a sealed bottle was sparged with N2 for 2 hours to remove oxygen, and then transferred into the glove box. Inside the glove box, 327 mg of [Ir(cod)Cl]2 and 716 mg of (S)-Xyl-BINAP solids were added to the dried and N2 purged pressure bottle. Next, 140 mL of inerted THF was added to the bottle, and the bottle was stirred overnight in the glove box. The filled pressure bottle with the pressure head was sealed and pressurized to ~10 psig inside the glove box and then transferred out of the glove box and weighed. Density of the catalyst solution was 0.89 g/mL, which was used to calculate the final volume, subsequent final molar concentration of the catalyst solution, and pumping rate in the continuous reactor. The catalyst solution was red/orange in color. This catalyst solution was stable for extended periods of time as long as it was kept inert. The catalyst solution was often prepared several days in advance of the continuous reaction. Preparation of catalyst solution for 15 kg campaign in 32 L continuous reactor. Catalyst solution was prepared in a glove box to avoid oxygen. The procedure was the same as written


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for preparing catalyst solution for the seasoning portion of the campaign. Table 17 shows the materials used. Table 17. Catalyst solution preparation for 15 kg campaign in 32 L continuous reactor. Material

Source

MW

Amount

[Ir(cod)Cl]2

UMICORE

671.71

4.5 g

(S)-Xyl-BINAP

Takasago

734.9

9.85 g

THF

Superior Solvents and Chemical, Indianapolis

72.11

0.430 L

A 1 liter pressure bottle was used instead of 300 mL. In the glove box, 4.5 g of [Ir(cod)Cl]2 and 9.85 g of (S)-Xyl-BINAP solids were added to the bottle, and then 430 mL of THF which had previously been sparged with N2. This solution was stirred in the glove box overnight. The bottle was sealed, pressurized with 10 psig nitrogen, removed from the glove box, and weighed. Preparation of substrate solution for seasoning run in 32 L continuous reactor. A 50 L glass flask was rinsed with THF and dried with N2 purge prior to addition of reagents. Materials charged to the 50 L flask are listed in Table 18. All feed preparations were done under nitrogen blanketing and purging. Table 18. Materials used for substrate solution for seasoning run in 32 L continuous reactor. Material

Source

MW

Amount

Ketal (1)

Vendor

293.36

572 g

Vendor

99.1

212 g

Aldrich

369.37

7.2 g

CSA

Acros

232.3

9.1 g

THF

Superior

72.11

40 L

Amino-tetrazole (2) TBAI

572g of the ketal (1) (light yellow powder), 212 g of the amino-tetrazole (2), 7.2 g of TBAI, and 9.1 g CSA were added to the 50 L flask equipped with overhead stirring and N2 sparge line. 4 L of THF was added at room temperature. The vessel was stirred until a homogeneous solution was obtained. Stirring was continued at room temperature with subsurface N2 sparging for 1 hour. After 1 h, the remaining 36 L of THF was added, and N2



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

sparging was continued for 2 more hours. Then this solution was transferred out of the 50L flask by vacuum and into a tared 16 gal pressure/vacuum vessel that had been previously cleaned and dried. The sealed 16 gallon vessel was pressurized to 15 psig and stored until the start of the continuous reactor. The density of the substrate feed solution was 0.934 g/mL. Preparation of substrate solution for 15 kg campaign in 32 L continuous reactor. The substrate solution for the 15 kg campaign was prepared in the same manner as the solution for the seasoning run. The same 50L glass vessel was used. Materials charged to the 50 L vessel are listed in Table 19. All feed preparations were done under nitrogen blanketing and purging. Table 19. Materials used for substrate solution for the 15 kg campaign in 32 L continuous reactor. Material

Source

MW

Amount

Ketal (1)

Vendor

293.36

15000 g

Amino-tetrazole (2)

Vendor

99.1

5574 g

TBAI

Aldrich

369.37

188.6 g

CSA

Acros

232.3

237.6 g

THF

Superior

72.11

105 L

Preparation of this solution was split up into 3 batches in the 50 L flask. Then after each of the feed batches were prepared, the solution was combined in a 55 gallon pressure/vacuum vessel and mixed. Each of the 3 sections was prepared as follows. 5000 g of the ketal (1) (light yellow powder), 1857 g of the amino tetrazole (2), 62.87 g of TBAI, and 79.19 g CSA were added to the 50 L flask equipped with overhead stirring and N2 sparge line. 35 L of THF was added at room temperature. The vessel was stirred until a homogeneous solution was obtained, and then stirring and N2 sparging was continued for 2 more hours. Then this solution was transferred out of the 50L flask by vacuum and into a tared 55 gallon portable pressure/vacuum vessel and held under N2 pressure. This was repeated 2 more times and all the reagent solution


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was stored in the 55 gallon portable vessel until the continuous campaign (all 3 sections mixed together). Operation of the 32 L horizontal pipes in series continuous reactor. ISCO high pressure syringe pumps were used for reagent and catalyst solution feeds to the continuous reactor. The reagent feed was pumped with dual 1000D ISCO pumps with automated switching valve package for uninterrupted continuous pumping mode. While one syringe pumped into the reactor, the other re-filled from the feed tank. The catalyst solution was pumped with a single 500D ISCO syringe pump. Flow rates were initially set for 12 hour τ and S:C 4000:1 in the reactor. Reagent feed (0.409 mol/L) pumping rate was 41.5 mL/min, and catalyst solution (0.0302 mol/L) pumping rate was 0.140 mL/min. Hydrogen feed gas was supplied by standard A size cylinders (197 SCF Hydrogen at STP, actual physical volume 40 L) which were initially pressurized to about 2100 psig. Hydrogen feed rate to the continuous reactor was controlled at about 0.06 mol/minute by automated cart 215, as described in the Supporting Information. Temperature was maintained at 115 °C for the continuous reactor by using feedback control for the amount of steam flowing through a stainless steel tubing coil submerged in the same oil bath as the reactor. Hydrogen and liquid flowed co-currently through the reactor, in the uphill direction through the larger diameter pipes and in the downhill direction through the smaller diameter connecting tubes, therefore gas τ was less than liquid τ because the gas bubbled past the liquid in the pipes. Operation of the online HPLC. All online HPLC data was collected using a Waters Corporation (Milford, Massachusetts, www.waters.com) PATrol UPLC Process Analysis System. A robust and reliable process-to-instrument interface is critical to generating reliable analytical data. The product solution was sampled downstream from the de-pressurizaiton zone.



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

The sampling zone was a fixed volume chamber between two pneumatically controlled valves, which were cycled on a timing schedule dependent on the reactor scale and flow rate. The upper valve opened allowing reaction material to enter and gravity drain into the lower portion of the chamber, including two ¼” Tees that were the process interface for online spectroscopic probes and the LC sampling interface. The sample remained in this zone for the length of a cycle, during which the process sampling module on the LC drew a sample from the reservoir. Prior to the next sample entering the zone, the remaining material in the zone exited via nitrogen push through a lower valve to a collection reservoir. The online LC interface used to collect all DARA reaction data was fluidically connected to the ¼” tee at the bottom of the sampling zone via a 1-m long, 0.030 inch i.d., PEEK tube. The PATrol HPLC system had a Process Sampling Module (PSM) that included sample draw capability using a peristaltic pump and sample dilution up to 100x. The sample line was connected at the bottom of the sample reservoir to minimize the possibility of nitrogen pulling into the sampling system. When the PATrol run sequence began, the peristaltic pump ran for 60 seconds at 4mL/min, delivering material to the process sampling valve onboard the LC. The sampling volume was determined based on leaving enough material in the reservoir so as not to pull it dry, while at the same time ensuring a representative sample was delivered to the dilution fluidics with the PATrol. To minimize carryover within LC fluidics, the volumes were kept intentionally small, while maximizing the number of turnovers per sample. The subsequent 100x dilution of process material was done by dynamically mixing the flowing sample stream with a diluent stream of 50/50 (V/V) acetonitrile/water. The diluted sample was then delivered to a isx-port, 2-position sampling valve with a 10 uL sample loop. The LC separation method was run at 40°C using a 10 mm ammonium acetate/methanol mobile phase system. The gradient and the analysis time, and thus the total analytical cycle time, varied


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based on the cycle time of the reactor cart (which was based on the volume of the reactor). Detection of species of interest was done using a UV flow cell at 220 nm wavelength. Supporting Information: The SI contains pictures of continuous reactors, constant temperature heating baths, and automated mini-plant carts for pressure and flow controls. The reactors in the pictures include 68 mL, 780 mL, and 32 L coiled tubes, 780 mL and 32 L horizontal pipes in series, and 10 mL, 410 mL, 11 L, and 360 L vertical pipes in series. The SI describes axial dispersion testing for continuous reactors and shows F-curve transitions and fitted axial dispersion numbers for each of the reactors used in this study. It describes the automated systems for control of reactor pressure, gas flow, de-pressurization at the reactor exit, and gas/liquid separation. Two pilot scale process control options are described in detail with schematic drawings, and the automation sequences are explained. Pressure versus time trends plots from the distributed control system are shown in figures, which help to explain how the systems worked. Option 1 does not use a back pressure regulator, while option 2 does. Option 1 can handle some solids in flow without clogging but has pressure oscillations. The research scale pressure and gas flow control system is also described in detail. Author Information: corresponding author *email: [email protected] Acknowledgements Takasago designed and developed the catalyst and ligand for the Ru catalyzed version of this DARA reaction, which was demonstrated in flow at research scale in the 400 ml vertical pipes in series reactor. Special acknowledgement is given to Tohru Yokozawa. Ed Deweese and



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Paul Milenbaugh constructed all of the continuous reactor systems, and Rick Spears operated them. Ed Plocharczyk consulted on design and development of the pipes in series reactors and eliminating surging. Jonathan Adler did all the kLa measurements in the tubing and pipes with water and air. These 5 people are all employees of D and M Continuous Solutions. Marv Hansen and Ryan Linder developed the Ir catalyzed DARA in toluene batch. Scott Frank and Radhe Vaid were the senior process chemists responsible for the DARA reaction scale up to manufacturing. V. Scott Sharp measured dissolved O2 with the on-line probe during the kLa characterizations. Phil Hoffman helped with design and development of the pressure and flow control systems that are described in detail in the Supporting Information. We thank Bret Huff for leading and sponsoring the continuous reaction design and development work at Eli Lilly and Company. References

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Design and Comparison of Tubular and Pipes-in-Series Continuous

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