10 replies to this topic

[chemtan]

PS: I have searched for information on below queries in Che Forums but could not find (although there are several student topics discussing this but do not cover below queries.)

Dear Friends,
Currently, I have to design LPG extraction plant which uses cryogenic process (Turbo expander). I have already devised a PFD (after going through several literatures) for it and in process of simulating this Process Flow.
The questions in my mind at the moment are:
1. The outlet temperature at Turbo expander is around -80 deg C, will CO2 cause icing at this temperature? How to prevent it or detect it during simulation? Gas composition contains CO2 ~ 3 mole%
2. As you can see in the attached PFD, I’m trying to remove C1+C2 in De-Ethanizer top and then C3+C4 in De-Butanizer top – is this fine? I mean having C2 in sales gas?
If someone has already designed a similar plant, please help by providing above information? Any comments on the PFD are also welcome and will be answered.
Attachments are:
a. Gas composition (XLS)
b. PFD in .pdf format

Attached Files

•  Gas_Composition.xls   16.5KB   163 downloads
•  TuroExpander_PFD.pdf   30.55KB   242 downloads

[JoeWong]

There is CO2 Feeze-out utility in HYSYS...

There are more improvements can be done on your present scheme for a better extraction. However, it may move into some patented design (which is non-disclosable...). Check out similar schemes in well-known licensor i.e. AP, FW, TECHNIP, etc.

[Zauberberg]

1. CO2 freezing is not an issue in LPG recovery plants. It becomes a concern in NGL extraction (C2+) and natural gas liquefaction plants. Search for technical materials within the Advanced Extraction Technologies (AET) website.

2. That's a conventional configuration (two-staged) for LPG and C5+ recovery. Usually, Sales Gas spec is defined by min/max heating values and max content of inert components, required cricondentherm/dew point etc. It all depends what specs have been given to you. Look at Ortloff's web page for excellent technical articles.

Best regards,

[chemtan]

Thanks for the so much required help in this regard. Please see below for some replies:

Answer #1:
Before I started the simulation, I went through the book:
Title: Oilfield Processing of Petroleum. Vol One: Natural Gas
Chapter # 14, page 317 says if CO2 contents are btw 0.5 - 1mole%, then there is a chance of CO2 freezing at outlet of turbo-expander OR in De-C1 top. The temperature sited in this book are around -90degC to -110degC.

when I simulated, the temperature at Turbo-Expander outlet is around -75degC... while De-C2 top has temperature of around -20degC so, I think my plant would not have this problem.


Answer#2:
Yes, I was given the LPG properties that I have to match at De-C4 top product.

Query #1:
I have simulated plant so that heat exchange takes place between De-C4 inlet stream and top vapors. The temperature difference btw inlet streams is aound 12degC and I have set the oultet temperature of vapors to a difference of 2degC only.

I could not use ambient air for this purpose as it is hotter than the vapors so devised this scheme to save cost.

The question that I cannot answer is that is it practically possible to construct a HX that will cool the stream to only 2degC?

[Zauberberg]

Here is one of the "off-the shelf" models attached for your reference.

I wouldn't expect such a high temperature for De-C3 tower top: it looks like you may be sacrificing precious recompression power in order to accomodate for separation at very low pressures. The diameter of your De-C3 tower (I'm still assuming we are speaking here about C3+ recovery and not C2+ recovery) will be extremely high as well. I would target something about 25-30 bar De-C3 tower pressure (close to expander outlet pressure).

As far as heat exchanger efficiency is concerned, looking for 3C temperature approach can be considered as a standard in modern cryogenic processing. Core-in-shell or printed circuit or BAHX can do this for you.

Good luck,

 LPG_Unit.zip   74.6KB   168 downloads

[chemtan]

Thanks for the file.

The design approach in your simulation is to obtain C1 + C2 @ tower top and LPG+other heavies @ bottom. I cannot employ this as there are more C5+ in my gas composition. So I had to top off the gas and then recover the LPG in tower downstream.

Having said that, it was still very helpful as it built my confidence on my design approach and I got sure that I'm on correct track - only tweaks needed.

However, I did learn one important thing from this simulation -> Converging the column on Product RVP specs. This seems like a more practical and direct method to get the required degree of LPG from tower. I used C5 mole% at top product and C4 mole% at bottom product which is good, but not like yours approach.

Thanks again.

Attached is the flow diagram frommy simulation (a PDF format), still in progress.

[chemtan]

Forgot to attach the flow diagram.. here it is.

BTW, there were few queries that I had in my mind regarding your simulation:

1. There is a cooler installed after cold box - is that a chiller?
2. why do you need "Set" Logical operation in your simulation?

Attached Files

•  LPG_Simulation.pdf   58.88KB   161 downloads

[JoeWong]

Yeap. Generally there is no problem for temperature of -75 degC as pure CO2 freezing point is about 78.5 degC.

Top of De-C4 will be a mixture of C3 & C4. The feed composition from field may change from time to time. So, the C3 & C4 level will change and DeC4 output will vary. Similarly to NGL absorber. Try to evaluate the operability of your system for lean, rich, high inert, high CO2, high ambient, etc.

You did not mentioned about the feed pressure. I would expect high pressure ( >60 barg). You may operate your DeC4 at 25-30 barg so that the temperature is high enough and you may employ air cooling. If you have low feed pressure, then it may force you into a refrigerant loop. You may to consider the make-up of refrigerant.

Don't quite understand you last question.
If you are talking about "stream temperature of 2 degC), yes. There is HX i.e. PHE, Spiral wond HX, S&T, etc will capable of cool the stream to 2 degC.

If you are talking about temperature approach of 2 degC, then still yes but it will be very challenging. Special surface need to used, costly, etc.





(i may have wrong understanding of what you said).

[chemtan]

Thanks Joe.

Actually, the design approach given by client states that design shall be based by considering expander inlet pressure of 800PSIG (54.4Bar).

Gas pressure at plant battery limit is given as 600PSIG (40.8Bar), it is then compressed to 800PSIG.

Client is not considering a refrigeration system for its plant and also does not want to have gas at pressure higher than 800 PSIG at Expander inlet. (I think I should communicate to them that gas pressure would need to be higher than 800PSIG if they dont want to use refrigeration).

And for 2degC, I meant a temperature approach of 2degC (Inlet Temp - Outlet Temp = 2degC). I think once I'm allowed to increase gas pressure at De-C4 inlet, this issue can be resolved because ambient air can be used for cooling.

[Padmakar Katre]

Dear,
The articles says that the temperature of (-) 90 to (-110) C but its for 0.5 to 1 mol% CO2 whereas in your case its 3 mol% so I think this should be carefully analysed.

[Padmakar Katre]

Dear,
The articles says that the temperature of (-) 90 to (-110) C but its for 0.5 to 1 mol% CO2 whereas in your case its 3 mol% so I think this should be carefully analysed.