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7

Heat Balance For A Steam Reformer Unit


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#251 PingPong

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Posted 29 May 2018 - 01:18 PM

Surely you must be familiar with the theory of heat transfer as part of your curriculum? Which textbook is used?

 

In any case there is a chapter about it in the Towler & Sinnott book that you mentioned a few days ago.

And in Coulson & Richardson's Chemical Engineering Volume 6 which is also written by Sinnott.

And in Perry.

And many others.



#252 MurtazaHakim

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Posted 31 May 2018 - 06:23 PM

We have started the preliminary simulation of the process. The simulation requires the temperature and pressure profile across the entire plant. Moreover the pressure of DM water at 30 oC, sat. BFW at 257 oC, sat. HPS at 257 oC and superheated HPS at 385 oC is required along with the pressure drop across the tube side and shell side of the heat exchangers employed between reformer-WGSR and WGSR-PSA beds. What values should we take for the above mentioned parameters ?



#253 PingPong

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Posted 02 June 2018 - 10:13 AM

sat HPS comes from the SD of which you know the pressure.

Take 1 bar for HPS superheat coil deltaP. Then HPS is split in process steam, which is sent to gas feed via a flow control valve, and export steam which is sent via a pressure control valve to B.L. at 41 bara (40 barg).

 

BFW pressure must be high enough to enter steam drum plus an allowance for pressure drop of preheat exch, lines, control valve, so take 60 bar for now. Exact value does not matter for simulation as long as it is high enough to avoid BFW vaporization.

 

Actual pressures depend on the actual designs of various equipment, so all you can do now is base profile on typical pressure drops.

 

Simulation pressures for gas streams do not really matter either except for reformer outlet (which you chose as 20 bara) as that affects equilibrium. So don't worry to much and make your own assumptions where necessary.

 

I suggest following based on your 20 bara reformer outlet, and typical equipment pressure drops as in industrial SMR designs:

 

20.0 bara Reformer outlet

19.7 bara WGSR inlet

19.4 bara WGSR outlet

18.0 bara Condensate separator and PSA inlet

17.5 bara PSA outlet

17.0 bara Hydrogen product to B.L. (downstream control valve)

 

17.5 bara Recycle hydrogen compressor inlet

27.0 bara Recycle hydrogen compressor outlet

 

36.0 bara Natural gas from B.L. (35 barg specified by you in the past)

27.0 bara Natural gas from flow control valve

26.5 bara Mixing point natural gas and recycle hydrogen

26.5 bara Inlet gas preheat coil

25.5 bara HDS inlet

24.5 bara ZnO outlet

24.0 bara Mixing point with process steam, upstream mixed feed preheat coil

23.0 bara Reformer inlet



#254 MurtazaHakim

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Posted 02 June 2018 - 12:18 PM

We are going forward with the Heat Exchanger design of WHB . The assumption of overall HTC value of 37 W/m2 oC  results in an approximate tube area of 30 m2 . Please look into the attached thumbnails and verify whether the parameters are correct or not.

Attached Files


Edited by MurtazaHakim, 02 June 2018 - 12:20 PM.


#255 PingPong

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Posted 02 June 2018 - 01:00 PM

I have no idea how you obtained that U = 37 W/m2.K but that is unrealistically low.



#256 MurtazaHakim

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Posted 02 June 2018 - 01:38 PM

A source from the internet suggested the approximate value 5-37 W/m2.K for gases in free convection. However another source suggests a value between 200-400 W/m2.K for gas at high pressure inside and liquid outside tubes. Taking U = 200 W/m2.K reduces the approximate tube area to 5.35 m2. What is the expected value of U for syngas at 20 bara pressure ?



#257 PingPong

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Posted 03 June 2018 - 04:47 AM

A source from the internet suggested the approximate value 5-37 W/m2.K for gases in free convection.

The syngas flow through the tubes of the WHB is not free convection, but forced convection.

another source suggests a value between 200-400 W/m2.K for gas at high pressure inside and liquid outside tubes.

That sounds more realistic.

Note moreover that in a WHB the film coefficient ho on the tube outside is determined by boiling, not simply by flowing liquid. Therefor it will be better than for liquid only, somewhere in the range of 5000 - 10000 W/m2.K. Exact value is difficult to calculate so you best take my word for it.

 

The film coefficient hi on the tubeside will be in the order of 500 - 1500 W/m2.K.

You could calculate hi by hand using the famous Nu = 0.023 * ................ equation.

It will be a rather small exchanger with one tube pass (special design).

For calculating hi use 28 tubes of 3/4 inch OD. That fits in an 8 inch shell when using a 1 inch (rotated) square pitch (see tube count tables in Perry). Tube wall thickness of say 2 mm.

 

Include fouling factors for shell and tube side, and the resistance of the tube wall, and you can calculate U.

Tube length is then resulting from calculated required A.

 

 

Your previous calculation of LMTD is not correct. 537 oC is way too high.

 

And as a WHB has only one tube pass, correction factor Ft is 1.00

Ft need only be calculated for multiple tube passes per shell.


Edited by PingPong, 03 June 2018 - 04:48 AM.


#258 MurtazaHakim

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Posted 05 June 2018 - 11:02 AM

In reformer section we considered 10 tubes of 4 inch diameter (114 mm OD and 87 mm ID) having total exposed surface area of around 17.5 m2 . What would be the typical density of Nickel catalyst filled in the reformer tubes ? The mixer is used for mixing the feed with recycle H2 at one stage and the process HPS at another stage. What is the typical construction of a mixer and what are the different components in an industrial mixer ? 



#259 PingPong

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Posted 05 June 2018 - 12:45 PM

Bulk density of reformer catalyst depends very much on height and diameter of the tablets (affects void fraction) and the number and size of the holes in the tablets (affects particle density). That all differs per vendor.

 

As far as I know bulk density of steam reforming cat varies in the range of 900 to 1100 kg/m3.

 

There are no mixers required for mixing of H2 and natgas or natgas and steam.

Just two pipes that are combined using a T-piece.



#260 MurtazaHakim

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Posted 07 June 2018 - 12:51 PM

Calculation of Nusselt number using Dittus-Boelter equation requires the Reynolds number and Prandtl number to be computed. Calculation of Reynolds number and Prandtl number in turn requires the physical properties of the fluid at the local bulk mean temperature. Furthermore for calculating the Reynolds number we need the velocity of the fluid and diameter of conduit. We can fix the diameter of conduit but how do we get the velocity of the fluid from the flowrate (115.8718 kmol/hr). How do we calculate  the local bulk mean temperature for calculation of physical properties of syngas ? The local bulk mean temperature cannot be just the arithmetic mean of the outlet and inlet temperature since the syngas experiences a huge temperature drop of around 500 degrees Celsius (inlet - 850 ⁰C , outlet - 353 ⁰C ).



#261 PingPong

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Posted 07 June 2018 - 01:19 PM

You know all the stream properties at inlet and outlet from your petrosim run, so calculating the volumetric flowrate of the syngas is a piece of cake (the run output probably also gives it).

 

I already gave you a starting point with number of tubes and diameter, so calculating the syngas velocity at inlet and outlet is again a piece of cake.

 

Do the calculation for Re, Pr, Nu, hi and U for the syngas inlet as well as the outlet and take average U.

Do not forget to include fouling factors in U calculation.


Edited by PingPong, 07 June 2018 - 01:21 PM.





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